Hydromethanation of a carbonaceous feedstock with improved carbon utilization and power generation

ABSTRACT

The present invention relates generally to processes for hydromethanating a carbonaceous feedstock in a hydromethanation reactor to a methane-enriched raw product stream, and more specifically to processing of solid char by-product removed from the hydromethanation reactor to improve the carbon utilization and thermal efficiency and economics of the overall process by co-producing electric power and steam from the by-product char in addition to the end-product pipeline quality substitute natural gas.

CROSS-REFERENCE TO RELATED APPLICATION

This application is related to U.S. application Ser. No. 16/171,858,filed Oct. 26, 2018, entitled HYDROMETHANATION OF A CARBONACEIOUSFEEDSTOCK WITH IMPROVED CARBON UTILIZATION), which is incorporated byreference herein for all purposes as if fully set forth below.

FIELD OF THE INVENTION

The present invention relates generally to processes forhydromethanating a carbonaceous feedstock in a hydromethanation reactorto a methane-enriched raw product stream, and more specifically to theprocessing of solid char by-product removed from the hydromethanationreactor to improve the carbon utilization and thermal efficiency of theoverall process, generate power for internal use and export, and therebylower the net costs of a pipeline quality substitute natural gas (SNG)end product.

BACKGROUND

In view of numerous factors such as higher energy prices andenvironmental concerns, the production of value-added products (such aspipeline-quality substitute natural gas, hydrogen, methanol, higherhydrocarbons, ammonia and electrical power) from lower-fuel-valuecarbonaceous feedstocks (such as petroleum coke, resids, asphaltenes,coal and biomass) is receiving renewed attention.

Such lower-fuel-value carbonaceous feedstocks can be gasified atelevated temperatures and pressures to produce a synthesis gas streamthat can subsequently be converted to such value-added products.

One advantageous gasification process is hydromethanation, in which thecarbonaceous feedstock is converted in a fluidized-bed hydromethanationreactor in the presence of a catalyst source and steam atmoderately-elevated temperatures and pressures to directly produce amethane-enriched synthesis gas stream raw product. This is distinct fromconventional gasification processes, such as those based on partialcombustion/oxidation of a carbon source at highly-elevated temperaturesand pressures (thermal gasification, typically non-catalytic), where asyngas (carbon monoxide+hydrogen) is the primary product (little or nomethane is directly produced), which can then be further processed toproduce methane (via catalytic methanation, see reaction (III) below) orany number of other higher hydrocarbon products.

Hydromethanation processes and the conversion/utilization of theresulting methane-rich synthesis gas stream to produce value-addedproducts are disclosed, for example, in U.S. Pat. Nos. 3,828,474,3,958,957, 3,998,607, 4,057,512, 4,092,125, 4,094,650, 4,204,843,4,243,639, 4,468,231, 4,500,323, 4,541,841, 4,551,155, 4,558,027,4,604,105, 4,617,027, 4,609,456, 5,017,282, 5,055,181, 6,187,465,6,790,430, 6,894,183, 6,955,695, US2003/0167691A1, US2006/0265953A1,US2007/0000177A1, US2007/0083072A1, US2007/0277437A1, US2009/0048476A1,US2009/0090056A1, US2009/0090055A1, US2009/0165383A1, US2009/0166588A1,US2009/0165379A1, US2009/0170968A1, US2009/0165380A1, US2009/0165381A1,US2009/0165361A1, US2009/0165382A1, US2009/0169449A1, US2009/0169448A1,US2009/0165376A1, US2009/0165384A1, US2009/0217582A1, US2009/0220406A1,US2009/0217590A1, US2009/0217586A1, US2009/0217588A1, US2009/0218424A1,US2009/0217589A1, US2009/0217575A1, US2009/0229182A1, US2009/0217587A1,US2009/0246120A1, US2009/0259080A1, US2009/0260287A1, US2009/0324458A1,US2009/0324459A1, US2009/0324460A1, US2009/0324461A1, US2009/0324462A1,US2010/0071235A1, US2010/0071262A1, US2010/0120926A1, US2010/0121125A1,US2010/0168494A1, US2010/0168495A1, US2010/0179232A1, US2010/0287835A1,US2010/0287836A1, US2010/0292350A1, US2011/0031439A1, US2011/0062012A1,US2011/0062721A1, US2011/0062722A1, US2011/0064648A1, US2011/0088896A1,US2011/0088897A1, US2011/0146978A1, US2011/0146979A1, US2011/0207002A1,US2011/0217602A1, US2011/0262323A1, US2012/0046510A1, US2012/0060417A1,US2012/0102836A1, US2012/0102837A1, US2012/0213680A1, US2012/0271072A1,US2012/0305848A1, US2013/0046124A1, US2013/0042824A1, US2013/0172640A1,US2014/0094636A1, WO2011/029278A1, WO2011/029282A1, WO2011/029283A1,WO2011/029284A1, WO2011/029285A1, WO2011/063608A1 and GB1599932, all ofwhich are hereby incorporated by reference. See also Chiaramonte et al,“Upgrade Coke by Gasification”, Hydrocarbon Processing, September 1982,pp. 255-257; and Kalina et al, “Exxon Catalytic Coal GasificationProcess Predevelopment Program, Final Report”, Exxon Research andEngineering Co., Baytown, TX, FE236924, December 1978, all of which arealso hereby incorporated by reference.

The hydromethanation of a carbon source typically involves fourtheoretically separate reactions:Steam carbon: C+H₂O→CO+H₂  (I)Water-gas shift: CO+H₂O→H₂+CO₂  (II)CO Methanation: CO+3H₂→CH₄+H₂O  (III)Hydro-gasification: 2H₂+C→CH₄  (IV)

In the hydromethanation reaction, the first three reactions (I-III)predominate to result in the following overall net reaction:2C+2H₂O→CH₄+CO₂  (V)

The overall hydromethanation reaction is essentially thermally balanced;however, due to process heat losses and other energy requirements (suchas required for evaporation of moisture entering the reactor with thefeedstock), some heat must be added to maintain the thermal balance.

Referring to FIG. 1, in one variation of the hydromethanation process,required carbon monoxide, hydrogen and heat energy can also at least inpart be generated in situ by feeding oxygen into the hydromethanationreactor. See, for example, previously incorporated US2010/0287835A1 andUS2011/0062721A1, as well as commonly-owned US2012/0046510A1,US2012/0060417A1, US2012/0102836A1, US2012/0102837A1, US2013/0046124A1,US2013/0042824A1, US2013/0172640A1 and US2014/0094636A1, andUS2010/0076235A1, which is hereby incorporated by reference. Thefollowing exothermic reactions will provide the heat energy to balancethe heat losses from the reactor and the energy requirement of the steamcarbon and other endothermic reactions:C+½O₂→CO  (VI)CO+½O₂→CO₂  (VII)C+O₂→CO₂  (VIII)H₂+½O₂→H₂O  (IX)

The result is a “direct” methane-enriched raw product gas stream alsocontaining substantial amounts of hydrogen, carbon monoxide and carbondioxide which can, for example, be directly utilized as a medium BTUenergy source, or can be processed to result in a variety ofhigher-value product streams such as pipeline-quality substitute naturalgas, high-purity hydrogen, methanol, ammonia, higher hydrocarbons,carbon dioxide (for enhanced oil recovery and industrial uses) andelectrical energy.

A char by-product stream is also produced in addition to themethane-enriched raw product gas stream. The solid char by-productcontains unreacted carbon, entrained hydromethanation catalyst and otherinorganic components of the carbonaceous feedstock. The by-product charmay contain 20 wt % or more carbon depending on the feedstockcomposition and hydromethanation conditions.

This by-product char is periodically or continuously removed from thehydromethanation reactor, and typically sent to a catalyst recovery andrecycle operation to improve economics and commercial viability of theoverall process. The nature of catalyst components associated with thechar extracted from a hydromethanation reactor and methods for theirrecovery are disclosed, for example, in previously incorporatedUS2007/0277437A1, US2009/0165383A1, US2009/0165382A1, US2009/0169449A1and US2009/0169448A1, as well as commonly-owned US2011/0262323A1 andUS2012/0213680A1. Catalyst recycle can be supplemented with makeupcatalyst as needed, such as disclosed in previously incorporatedUS2009/0165384A1.

Catalyst recovery, and particularly catalyst recovery from high-ash charproducts, is often difficult to perform economically due to severalfactors, including the volume of by-product char material produced.

In addition, the residue of the char after catalyst recovery stillcontains a significant amount of carbon that is unconverted and iseffectively removed from the process.

It would, therefore, be desirable to find a way to more completely andefficiently utilize the carbon in the char by-product from thehydromethanation reactor while reducing the volume of material that isultimately sent for catalyst recovery.

And further, in some cases it may be desirable to produce additionalelectric power and steam from the char by-product.

SUMMARY OF THE INVENTION

The optimum carbon conversion for a catalytic gasification process isgenerally restricted to be less than 100% to enable the rate of thesimultaneous gasification, shift and methanation reactions to be attheir maximum values. Typically, the organic carbon content of the bedis maintained at a value that restricts the carbon conversion to85%-95%. The high efficiency of catalytic gasification to producesubstitute natural gas relative to noncatalytic gasification can beachieved despite the lower carbon conversion.

The present invention discloses alternative embodiments to utilize thecarbon that is deliberately left unconverted in the hydromethanationreactor during a catalytic gasification process. The unused carbon inthe hydromethanation reactor represents a source of energy, eitherdirectly as heat or indirectly as synthesis gas, that can potentiallyfurther increase the efficiency of the process and improve theeconomics. In general, the unconverted char may be directly reacted withair or oxygen and steam in an oxidation reactor to produce either asynthesis gas (H₂, CO, CO₂) with little or no CH₄ or a combustionproduct (CO₂, H₂O) that can be utilized to produce additional steam andelectric power. The process in accordance with the present invention maybe useful because it improves carbon utilization in the process,improves the overall thermal efficiency, creates steam and electricpower, and reduces the volume of char material processed in catalystrecovery (which allows the use of reduced equipment size), therebylowering both operating and capital costs of the operation.

In particular, the present invention provides a process for generating,from a non-gaseous carbonaceous material and a hydromethanationcatalyst, (1) a fines-cleaned methane-enriched raw product gas stream,(2) an oxidation reactor product gas and (3) an oxidation reactor ashproduct stream, the process comprising the steps of:

a) preparing a carbonaceous feedstock from the non-gaseous carbonaceousmaterial;

b) introducing the carbonaceous feedstock, the hydromethanationcatalyst, steam and oxygen into a hydromethanation reactor, thehydromethanation reactor comprising a fluidized bed, a disengagementzone above the fluidized bed, and a gas mixing zone below the fluidizedbed;

c) reacting the carbonaceous feedstock in the hydromethanation reactorin the presence of carbon monoxide, hydrogen, steam and hydromethanationcatalyst, and at an operating temperature from about 400° F. (about 205°C.) up to about 1500° F. (about 816° C.), and an operating pressure ofat least about 250 psig (about 1825 kPa), to produce a methane-enrichedraw product gas, heat energy and a by-product char;

d) withdrawing a stream of methane-enriched raw product gas from thehydromethanation reactor as the methane-enriched raw product gas stream,wherein the methane-enriched raw product gas stream comprises methane,carbon monoxide, hydrogen, carbon dioxide, hydrogen sulfide, steam, andentrained solids;

e) removing a substantial portion of the entrained solids from themethane-enriched raw product gas stream to generate a solids-depleted,methane-enriched raw product gas stream and a recovered primary solidsstream;

f) removing a substantial portion of any fines from the solids-depleted,methane-enriched raw product gas stream to generate the fines-cleanedmethane-enriched raw product gas stream and a recovered secondary finesstream;

g) withdrawing a stream of by-product char from the hydromethanationreactor as the hydromethanation reactor char product stream, wherein thehydromethanation reactor char product stream comprises a carbon contentand entrained hydromethanation catalyst;

h) extracting a portion of the entrained catalyst from thehydromethanation reactor char product stream by the steps of:

(i) feeding all or a portion of the hydromethanation reactor charproduct stream to a catalyst recovery unit;

(ii) withdrawing a stream of catalyst-depleted char from the catalystrecovery unit as the washed char product stream; and

(iii) withdrawing a stream of liberated hydromethanation catalyst fromthe catalyst recovery unit as the recovered hydromethanation catalyststream;

i) feeding into an oxidation reactor (1) all or a portion of the washedchar product stream, and (2) an oxygen-containing gas stream;

j) reacting at least a portion of the carbon content of the washed charproduct stream with the oxygen-containing gas stream in the oxidationreactor to produce (1) the oxidation reactor ash product stream, and (2)the oxidation reactor product gas comprising steam and carbon dioxide;and

k) withdrawing the oxidation reactor product gas from the oxidationreactor.

In addition, the process may further comprise the steps of: (A)operating the oxidation reactor at a pressure greater than 791 kPa (100psig); (B) (i) when the oxidation reactor product gas further comprisescarbon monoxide and hydrogen, feeding all or a portion of the oxidationreactor product gas into a gas turbine combined cycle system to generatepower and steam; or (ii) when carbon monoxide and hydrogen are notpresent in the oxidation reactor product gas, feeding all or a portionof the oxidation reactor product gas into a gas expander and a heatrecovery steam generator to generate power and steam; and (C) cooling orquenching the oxidation reactor ash product stream. In addition, thesteps of (D) converting all or a portion of the steam generated in step(B) to power in a steam-turbine generator, and (E) exporting any unusedsteam are also contemplated.

The process may further comprise the steps of: (A) operating theoxidation reactor at a pressure less than 400 kPa (44 psig); (B) (i)when the oxidation reactor product gas further comprises carbon monoxideand hydrogen, feeding all or a portion of the oxidation reactor productgas to a heat recovery steam generator equipped with an auxiliary burnerto generate steam; or (ii) when carbon monoxide and hydrogen are notpresent in the oxidation reactor product gas, feeding all or a portionof the oxidation reactor product gas to a heat recovery steam generatorto generate steam; and (C) cooling or quenching the oxidation reactorash product stream. The steps of (D) converting all or a portion of thesteam generated in step (B) to power in a steam-turbine generator, and(E) exporting any unused steam are also contemplated. The operatingpressure in the hydromethanation reactor ranges from at least about 250psig (about 1825 kPa) up to about 1200 psig (about 8375 kPa). Theoxidation reactor may be a fluidized-bed oxidation reactor. Thehydromethanation catalyst may comprise an alkali metal, which may bepotassium. The hydromethanation catalyst may comprise at least a portionof the recovered hydromethanation catalyst stream.

The process may further comprise the steps of: (f1) feeding therecovered primary solids stream into the hydromethanation reactor; and(f2) feeding all or a portion of the recovered secondary fines streaminto the catalyst recovery unit.

The process may further comprise the steps of:

1) introducing the fines-cleaned methane-enriched raw product gas streaminto a heat exchanger unit to remove heat energy and generate a cooledmethane-enriched raw product stream;

m) steam shifting a portion of the carbon monoxide in the cooledmethane-enriched raw product stream in a shift reactor system togenerate a hydrogen-enriched raw product gas stream with a molar ratioof hydrogen to carbon monoxide of close to 3;

n) dehydrating the hydrogen-enriched raw product gas stream in alow-temperature gas cooling system, to generate a dry raw gas stream;and

o) removing a substantial portion of the carbon dioxide and asubstantial portion of the hydrogen sulfide from the dry raw gas streamin an acid gas removal unit to produce a sweetened gas stream comprisinga substantial portion of the hydrogen, carbon monoxide and methane fromthe dry raw gas stream.

The process may further comprise the steps of:

p) reacting the carbon monoxide and hydrogen in the sweetened gas streamin a methanation system in the presence of a methanation catalyst toproduce heat energy and a pipeline quality substitute natural gasstream;

q) recovering the heat energy from the catalytic methanation; and

r) utilizing at least a portion of the recovered heat energy to generateand superheat a steam stream; and may have a steam demand and a powerdemand that are met by internal energy integration such that the processrequires no net import of steam or power.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 is a general diagram of the basic hydromethanation process toconvert a hydrocarbonaceous feedstock to a fines-cleanedmethane-enriched raw product gas stream and a washed char productstream.

FIGS. 2 and 3 are general diagrams of processes for converting thewashed char product stream generated by the basic hydromethanationprocess to an oxidation reactor product gas in an oxidation reactor andutilizing the oxidation reactor product gas to make steam and electricpower, thereby maximizing the overall carbon conversion and overallthermal efficiency.

FIG. 2 is a general diagram of an embodiment wherein the oxidationreactor is operated at a high pressure greater than 791 kPa (100 psig)using oxygen or air as the oxidant and steam as a reactant to produce anoxidation gas that is further processed in a gas-turbine combined cycleto produce electric power and steam for export.

FIG. 3 is a general diagram of an alternative embodiment wherein theoxidation reactor uses air as an oxidant and optionally steam as areactant to produce an oxidation gas that is expanded to produceelectric power if at a high pressure and further processed in a heatrecovery steam generation system to produce electric power and steam forexport.

FIG. 4 is a diagram of a representative embodiment for the furtherprocessing of the fines-cleaned methane-enriched raw product stream tomake a pipeline-quality substitute natural gas (SNG).

DEFINITION OF TERMS

The present invention relates to processes for converting a non-gaseouscarbonaceous material ultimately into one or more value-added gaseousproducts.

In the context of the present description, all publications, patentapplications, patents and other references mentioned herein, if nototherwise indicated, are explicitly incorporated by reference herein intheir entirety for all purposes as if fully set forth.

Unless otherwise defined, all technical and scientific terms used hereinhave the same meaning as commonly understood by one of ordinary skill inthe art to which this disclosure belongs. In case of conflict, thepresent specification, including definitions, will control.

Except where expressly noted, trademarks are shown in upper case.

Unless stated otherwise, all percentages, parts, ratios, etc., are byweight.

Unless stated otherwise, pressures expressed in psi units are gauge, andpressures expressed in kPa units are absolute. Pressure differences,however, are expressed as absolute (for example, pressure 1 is 25 psihigher than pressure 2).

When an amount, concentration, or other value or parameter is given as arange, or a list of upper and lower values, this is to be understood asspecifically disclosing all ranges formed from any pair of any upper andlower range limits, regardless of whether ranges are separatelydisclosed. Where a range of numerical values is recited herein, unlessotherwise stated, the range is intended to include the endpointsthereof, and all integers and fractions within the range. It is notintended that the scope of the present disclosure be limited to thespecific values recited when defining a range.

When the term “about” is used in describing a value or an end-point of arange, the disclosure should be understood to include the specific valueor end-point referred to.

As used herein, the terms “comprises,” “comprising,” “includes,”“including,” “has,” “having” or any other variation thereof, areintended to cover a non-exclusive inclusion. For example, a process,method, article, or apparatus that comprises a list of elements is notnecessarily limited to only those elements but can include otherelements not expressly listed or inherent to such process, method,article, or apparatus.

Further, unless expressly stated to the contrary, “or” and “and/or”refers to an inclusive and not to an exclusive. For example, a conditionA or B, or A and/or B, is satisfied by any one of the following: A istrue (or present) and B is false (or not present), A is false (or notpresent) and B is true (or present), and both A and B are true (orpresent).

The use of “a” or “an” to describe the various elements and componentsherein is merely for convenience and to give a general sense of thedisclosure. This description should be read to include one or at leastone, and the singular also includes the plural unless it is obvious thatit is meant otherwise.

The term “substantial”, as used herein, unless otherwise defined herein,means that greater than about 90% of the referenced material, preferablygreater than about 95% of the referenced material, and more preferablygreater than about 97% of the referenced material. If not specified, thepercent is on a molar basis when reference is made to a molecule (suchas methane, carbon dioxide, carbon monoxide and hydrogen sulfide), andotherwise is on a weight basis (such as for entrained solids).

The term “predominant portion”, as used herein, unless otherwise definedherein, means that greater than 50% of the referenced material. If notspecified, the percent is on a molar basis when reference is made to amolecule (such as hydrogen, methane, carbon dioxide, carbon monoxide andhydrogen sulfide), and otherwise is on a weight basis (such as forentrained solids).

The term “depleted” is synonymous with reduced from originally present.For example, removing a substantial portion of a material from a streamwould produce a material-depleted stream that is substantially depletedof that material. Conversely, the term “enriched” is synonymous withgreater than originally present.

The term “carbonaceous” as used herein is synonymous with hydrocarbon.

The term “carbonaceous material” as used herein is a material containingorganic hydrocarbon content. Carbonaceous materials can be classified asbiomass or non-biomass materials as defined herein.

The term “biomass” as used herein refers to carbonaceous materialsderived from recently (for example, within the past 100 years) livingorganisms, including plant-based biomass and animal-based biomass. Forclarification, biomass does not include fossil-based carbonaceousmaterials, such as coal. For example, see previously incorporatedUS2009/0217575A1, US2009/0229182A1 and US2009/0217587A1.

The term “plant-based biomass” as used herein means materials derivedfrom green plants, crops, algae, and trees, such as, but not limited to,sweet sorghum, bagasse, sugarcane, bamboo, hybrid poplar, hybrid willow,albizia trees, eucalyptus, alfalfa, clover, oil palm, switchgrass,sudangrass, millet, jatropha, and miscanthus (e.g., Miscanthus xgiganteus). Biomass further include wastes from agriculturalcultivation, processing, and/or degradation such as corn cobs and husks,corn stover, straw, nut shells, vegetable oils, canola oil, rapeseedoil, biodiesels, tree bark, wood chips, sawdust, and yard wastes.

The term “animal-based biomass” as used herein means wastes generatedfrom animal cultivation and/or utilization. For example, biomassincludes, but is not limited to, wastes from livestock cultivation andprocessing such as animal manure, guano, poultry litter, animal fats,and municipal solid wastes (e.g., sewage).

The term “non-biomass”, as used herein, means those carbonaceousmaterials which are not encompassed by the term “biomass” as definedherein. For example, non-biomass include, but is not limited to,anthracite, bituminous coal, sub-bituminous coal, lignite, petroleumcoke, asphaltenes, liquid petroleum residues or mixtures thereof. Forexample, see US2009/0166588A1, US2009/0165379A1, US2009/0165380A1,US2009/0165361A1, US2009/0217590A1 and US2009/0217586A1.

“Liquid heavy hydrocarbon materials” are viscous liquid or semi-solidmaterials that are flowable at ambient conditions or can be madeflowable at elevated temperature conditions. These materials aretypically the residue from the processing of hydrocarbon materials suchas crude oil. For example, the first step in the refining of crude oilis normally a distillation to separate the complex mixture ofhydrocarbons into fractions of differing volatility. A typicalfirst-step distillation requires heating at atmospheric pressure tovaporize as much of the hydrocarbon content as possible withoutexceeding an actual temperature of about 650° F., since highertemperatures may lead to thermal decomposition. The fraction which isnot distilled at atmospheric pressure is commonly referred to as“atmospheric petroleum residue”. The fraction may be further distilledunder vacuum, such that an actual temperature of up to about 650° F. canvaporize even more material. The remaining undistillable liquid isreferred to as “vacuum petroleum residue”. Both atmospheric petroleumresidue and vacuum petroleum residue are considered liquid heavyhydrocarbon materials for the purposes of the present process.

Non-limiting examples of liquid heavy hydrocarbon materials includevacuum resids; atmospheric resids; heavy and reduced petroleum crudeoils; pitch, asphalt and bitumen (naturally occurring as well asresulting from petroleum refining processes); tar sand oil; shale oil;bottoms from catalytic cracking processes; coal liquefaction bottoms;and other hydrocarbon feedstreams containing significant amounts ofheavy or viscous materials such as petroleum wax fractions.

The term “asphaltene” as used herein is an aromatic carbonaceous solidat room temperature, and can be derived, for example, from theprocessing of crude oil and crude oil tar sands. Asphaltenes may also beconsidered liquid heavy hydrocarbon feedstocks.

The liquid heavy hydrocarbon materials may inherently contain minoramounts of solid carbonaceous materials, such as petroleum coke and/orsolid asphaltenes, that are generally dispersed within the liquid heavyhydrocarbon matrix, and that remain solid at the elevated temperatureconditions utilized as the feed conditions for the present process.

The terms “petroleum coke” and “petcoke” as used herein include both (i)the solid thermal decomposition product of high-boiling hydrocarbonfractions obtained in petroleum processing (heavy residues—“residpetcoke”); and (ii) the solid thermal decomposition product ofprocessing tar sands (bituminous sands or oil sands—“tar sandspetcoke”). Such carbonization products include, for example, green,calcined, needle and fluidized bed petcoke.

Resid petcoke can also be derived from a crude oil, for example, bycoking processes used for upgrading heavy-gravity residual crude oil(such as a liquid petroleum residue), which petcoke contains ash as aminor component, typically about 1.0 wt % or less, and more typicallyabout 0.5 wt % of less, based on the weight of the coke. Typically, theash in such lower-ash cokes predominantly comprises metals such asnickel and vanadium.

Tar sands petcoke can be derived from an oil sand, for example, bycoking processes used for upgrading oil sand. Tar sands petcoke containsash as a minor component, typically in the range of about 2 wt % toabout 12 wt %, and more typically in the range of about 4 wt % to about12 wt %, based on the overall weight of the tar sands petcoke.Typically, the ash in such higher-ash cokes predominantly comprisesmaterials such as silica and/or alumina.

Petroleum coke can comprise at least about 70 wt % carbon, at leastabout 80 wt % carbon, or at least about 90 wt % carbon, based on thetotal weight of the petroleum coke. Typically, the petroleum cokecomprises less than about 20 wt % inorganic compounds, based on theweight of the petroleum coke.

The term “coal” as used herein means peat, lignite, sub-bituminous coal,bituminous coal, anthracite, or mixtures thereof. In certainembodiments, the coal has a carbon content of less than about 85%, orless than about 80%, or less than about 75%, or less than about 70%, orless than about 65%, or less than about 60%, or less than about 55%, orless than about 50% by weight, based on the total coal weight. In otherembodiments, the coal has a carbon content ranging up to about 85%, orup to about 80%, or up to about 75% by weight, based on the total coalweight. Examples of useful coal include, but are not limited to,Illinois #6, Pittsburgh #8, Beulah (ND), Utah Blind Canyon, and PowderRiver Basin (PRB) coals. Anthracite, bituminous coal, sub-bituminouscoal, and lignite coal may contain about 10 wt %, from about 5 to about7 wt %, from about 4 to about 8 wt %, and from about 9 to about 11 wt %,ash by total weight of the coal on a dry basis, respectively. However,the ash content of any particular coal source will depend on the rankand source of the coal, as is familiar to those skilled in the art. See,for example, “Coal Data: A Reference”, Energy InformationAdministration, Office of Coal, Nuclear, Electric and Alternate Fuels,U.S. Department of Energy, DOE/EIA-0064(93), February 1995.

The ash produced from combustion of a coal typically comprises both afly ash and a bottom ash, as is familiar to those skilled in the art.The fly ash from a bituminous coal can comprise from about 20 to about60 wt % silica and from about 5 to about 35 wt % alumina, based on thetotal weight of the fly ash. The fly ash from a sub-bituminous coal cancomprise from about 40 to about 60 wt % silica and from about 20 toabout 30 wt % alumina, based on the total weight of the fly ash. The flyash from a lignite coal can comprise from about 15 to about 45 wt %silica and from about 20 to about 25 wt % alumina, based on the totalweight of the fly ash. See, for example, Meyers, et al. “Fly Ash. AHighway Construction Material,” Federal Highway Administration, ReportNo. FHWA-IP-76-16, Washington, D.C., 1976.

The bottom ash from a bituminous coal can comprise from about 40 toabout 60 wt % silica and from about 20 to about 30 wt % alumina, basedon the total weight of the bottom ash. The bottom ash from asub-bituminous coal can comprise from about 40 to about 50 wt % silicaand from about 15 to about 25 wt % alumina, based on the total weight ofthe bottom ash. The bottom ash from a lignite coal can comprise fromabout 30 to about 80 wt % silica and from about 10 to about 20 wt %alumina, based on the total weight of the bottom ash. See, for example,Moulton, Lyle K. “Bottom Ash and Boiler Slag,” Proceedings of the ThirdInternational Ash Utilization Symposium, U.S. Bureau of Mines,Information Circular No. 8640, Washington, D.C., 1973.

A material such as methane can be biomass or non-biomass under the abovedefinitions depending on its source of origin.

A “non-gaseous” material is substantially a liquid, semi-solid, solid ormixture at ambient conditions. For example, coal, petcoke, asphalteneand liquid petroleum residue are non-gaseous materials, while methaneand natural gas are gaseous materials.

The term “unit” refers to a unit operation. When more than one “unit” isdescribed as being present, those units are operated in a parallelfashion unless otherwise stated. A single “unit”, however, may comprisemore than one of the units in series, or in parallel, depending on thecontext. For example, an acid gas removal unit may comprise a hydrogensulfide removal unit followed in series by a carbon dioxide removalunit. As another example, a contaminant removal unit may comprise afirst removal unit for a first contaminant followed in series by asecond removal unit for a second contaminant. As yet another example, acompressor may comprise a first compressor to compress a stream to afirst pressure, followed in series by a second compressor to furthercompress the stream to a second (higher) pressure.

The term “a portion of the carbonaceous feedstock” refers to carboncontent of unreacted feedstock as well as partially reacted feedstock,as well as other components that may be derived in whole or part fromthe carbonaceous feedstock (such as carbon monoxide, hydrogen andmethane). For example, “a portion of the carbonaceous feedstock”includes carbon content that may be present in by-product char andrecycled fines, which char is ultimately derived from the originalcarbonaceous feedstock.

The term “char” is the combustible residue remaining after thedestructive distillation or partial gasification or partial oxidation ofcoal, petcoke or any other carbonaceous material. Besides thecombustible residue, char will contain any mineral matter in thecarbonaceous material, including non-volatile portions of catalyst.

“Catalyst-depleted char” is a char material which has undergone acatalyst removal process to recover and reuse soluble catalystcomponents.

“Carbon-depleted char” is char which has undergone a secondaryprocessing step such as oxidation or gasification or both to convert thecarbon residue to gaseous compounds.

The term “entrained catalyst” as used herein means chemical compoundscomprising the catalytically active portion of the hydromethanationcatalyst, e.g., alkali metal compounds present in the char by-product.For example, “entrained catalyst” can include, but is not limited to,soluble alkali metal compounds (such as alkali metal carbonates, alkalimetal hydroxides and alkali metal oxides) and/or insoluble alkalicompounds (such as alkali metal aluminosilicates). The nature ofcatalyst components associated with the char extracted are discussed,for example, in previously incorporated US2007/0277437A1,US2009/0165383A1, US2009/0165382A1, US2009/0169449A1 andUS2009/0169448A1.

The term “superheated steam” refers to a steam stream that isnon-condensing under the conditions utilized, as is commonly understoodby persons of ordinary skill in the relevant art.

The term “steam demand” refers to the amount of steam that must be addedto the hydromethanation reactor via the gas feed streams to thehydromethanation reactor. Steam is consumed in the hydromethanationreaction and some steam must be added to the hydromethanation reactor.The theoretical consumption of steam is two moles for every two moles ofcarbon in the feed to produce one mole of methane and one mole of carbondioxide (see equation (V)). In actual practice, the steam consumption isnot perfectly efficient and steam is withdrawn with the product gases;therefore, a greater than theoretical amount of steam needs to be addedto the hydromethanation reactor, which added amount is the “steamdemand”. Steam can be added, for example, via the steam stream (12) andthe oxygen stream (15), which are typically combined prior tointroduction into the hydromethanation reactor. The amount of steam tobe added (and the source) is discussed in further detail below. Steamgenerated in situ from the carbonaceous feedstock (e.g., fromvaporization of any moisture content of the carbonaceous feedstock, orfrom an oxidation reaction with hydrogen, methane and/or otherhydrocarbons present in or generated from the carbonaceous feedstock)can assist in providing steam; however, it should be noted that anysteam generated in situ or fed into the hydromethanation reactor at atemperature lower than the operating temperature within thehydromethanation reactor (the hydromethanation reaction temperature)will have an impact on the “heat demand” for the hydromethanationreaction.

The term “heat demand” refers to the amount of heat energy that must beadded to the hydromethanation reactor (for example, via steam stream(12) and/or generated in situ (for example, via a combustion/oxidationreaction with supplied oxygen stream (15) as discussed below) to keepthe reaction of carbonaceous feedstock in the hydromethanation reactorin substantial thermal balance, as further detailed below.

The term “power demand” refers to the amount of power that must be usedto operate the processes.

The term “substitute natural gas” or “SNG” refers to a methane-richproduct gas, typically with a higher heating value greater than 950BTU/scf, that meets all the specifications for natural gas prescribed bynatural gas merchants or pipeline carriers or operators.

Although methods and materials similar or equivalent to those describedherein can be used in the practice or testing of the present disclosure,suitable methods and materials are described herein. The materials,methods, and examples herein are thus illustrative only and, except asspecifically stated, are not intended to be limiting.

DETAILED INVENTION DESCRIPTION

Basic Process

The basic hydromethanation reactor (HMR) process is shown in FIG. 1.Referring to FIG. 1, the non-gaseous carbonaceous material (10) isprocessed in a feedstock preparation unit (100) to generate acarbonaceous feedstock (32) which is fed to a catalyst application unit(350) where hydromethanation catalyst (31) is applied to generate acatalyzed carbonaceous feedstock (31+32). The application methods caninclude mechanical mixing devices to disperse the catalyst solution overthe solid feed particles and thermal dryers to achieve the preferredmoisture content for the catalyzed carbonaceous feedstocks (31+32).

The feedstock preparation unit (100) includes coal or coke pulverizationmachines to achieve a pre-determined optimal size distribution whichlargely depends on the carbonaceous mechanical and chemical properties.In some cases, pelletization and/or briquetting machines are included toconsolidate fines to maximize the utilization of all solid feedstockmaterials. Further details are provided below.

The hydromethanation catalyst (31) will typically comprise a recoveredhydromethanation catalyst stream (57) recovered from hydromethanationreactor char product stream (54) and recovered secondary fines (66), anda make-up catalyst from a make-up catalyst stream (56). Further detailsare provided below.

The catalyzed carbonaceous feedstock (31+32) is fed into ahydromethanation reactor (200) along with steam stream (12) and oxygenstream (15). The location at which the catalyzed carbonaceous feedstock(31+32) is fed may vary. For example, it may be fed into fluidized bed(202) in lower portion (202 a), just above the hydromethanation reactorfluidizing gas distributor plate (208), or into the bottom of upperportion (202 b).

Steam streams (12) and (12 a) are provided by a steam source such assteam distribution system (11), which desirably utilizes process heatrecovery (e.g., heat energy recovery from the hot raw product gas andother process sources) such that the process is steam integrated andsteam sufficient. Oxygen stream (15) and second oxygen stream (15 a),which splits off from oxygen stream (15), is supplied by an airseparation unit (14).

The steam stream (12) and oxygen stream (15) may be a single feed streamwhich comprises, or multiple feed streams which comprise, in combinationwith the in situ generation of heat energy and syngas, steam and heatenergy, as required to at least substantially satisfy, or at leastsatisfy, steam and heat demands of the hydromethanation reaction thattakes place in hydromethanation reactor (200).

In the hydromethanation reactor (200), (i) a portion of the carbonaceousfeedstock, steam, hydrogen and carbon monoxide react in the presence ofthe hydromethanation catalyst to generate a methane-enriched raw productgas (the hydromethanation reaction), and (ii) a portion of thecarbonaceous feedstock reacts in the presence of steam and oxygen togenerate heat energy and typically carbon monoxide, hydrogen and carbondioxide (combustion/oxidation reaction). The generated methane-enrichedraw product gas is withdrawn from the hydromethanation reactor (200) asa methane-enriched raw product gas stream (50). The withdrawnmethane-enriched raw product gas stream (50) typically comprises atleast methane, carbon monoxide, carbon dioxide, hydrogen, hydrogensulfide, steam, and entrained solids.

The hydromethanation reactor (200) comprises a fluidized bed (202)having an upper portion (202 b) above a lower portion (202 a) and adisengagement zone (204) above the fluidized bed. Hydromethanationreactor (200) also typically comprises a gas mixing zone (206) below thefluidized-bed (202), with the two sections typically being separated bya hydromethanation reactor fluidizing gas distributor plate (208) orsimilar divider (for example, an array of sparger pipes). Oxygen (15) ismixed with the high-pressure, superheated steam (12), and the mixtureintroduced into the gas mixing zone (206), into the lower portion (202a) of the fluidized bed (202) via the gas mixing zone (206), into thefluidized bed (202) at other locations, or into a combination thereof,via the hydromethanation reactor fluidizing gas distributor plate (208)or similar divider. Desirably, oxygen is fed into the lower portion ofthe fluidized bed. Without being bound by any theory, thehydromethanation reaction predominates in upper portion (202 b), and anoxidation reaction with oxygen predominates in lower portion (202 a). Itis believed that there is no specific defined boundary between the twoportions, but rather there is a transition as oxygen is consumed (andheat energy and syngas are generated) in lower portion (202 a). It isalso believed that oxygen consumption is rapid under the conditionspresent in hydromethanation reactor (200).

At least a portion of the carbonaceous feedstock in lower portion (202a) of fluidized bed (202) will react with oxygen from oxygen stream (15)to generate heat energy, and hydrogen and carbon monoxide (syngas). Thisincludes the reaction of solid carbon from unreacted (fresh) feedstock,partially reacted feedstock (such as char and recycled fines), as wellas gases (carbon monoxide, hydrogen, methane and higher hydrocarbons)that may be generated from or carried with the feedstock and recyclefines in lower portion (202 a). Generally, some water (steam) may beproduced, as well as other by-products such as carbon dioxide dependingon the extent of combustion/oxidation and the water gas shift reaction.As indicated above, in hydromethanation reactor (200) (predominantly inupper portion (202 b) of fluidized bed (202)) the carbonaceousfeedstock, steam, hydrogen and carbon monoxide react in the presence ofthe hydromethanation catalyst to generate a methane-enriched rawproduct, which is ultimately withdrawn as a methane-enriched raw productgas stream (50) from the hydromethanation reactor (200).

The reactions of the carbonaceous feedstock in fluidized bed (202) alsoresult in a by-product char comprising unreacted carbon as well asnon-carbon content from the carbonaceous feedstock (includinghydromethanation catalyst) as described in further detail below. Toprevent buildup of the residue in the hydromethanation reactor (200), asolid purge of hydromethanation reactor char product stream (54) isroutinely withdrawn (periodically or continuously) via a char withdrawalline (210). The char product stream (54) comprises a carbon content andentrained hydromethanation catalyst.

Char may also be withdrawn from the hydromethanation reactor at otherlocations such as from the top of fluidized bed (202), at any placewithin upper portion (202 b) and/or lower portion (202 a) of fluidizedbed (202), and/or at or just below hydromethanation reactor fluidizinggas distributor plate (208). For example, in one embodiment as disclosedin previously incorporated US2012/0102836A1, carbonaceous feedstock (32)(or catalyzed carbonaceous feedstock (31+32)) is fed into lower portion(202 a) of fluidized bed (202). Because catalyzed carbonaceous feedstock(31+32) is introduced into lower portion (202 a) of fluidized bed (202),at least one char withdrawal line (210) will typically be located at apoint such that by-product char is withdrawn from fluidized bed (202) atone or more points above the feed location of catalyzed carbonaceousfeedstock (31+32), typically from upper portion (202 b) of fluidized bed(202).

Particles too large to be fluidized in fluidized-bed section (202), forexample, large-particle by-product char and non-fluidizableagglomerates, are generally collected in lower portion (202 a) offluidized bed (202), as well as in gas mixing zone (206). Such particleswill typically comprise a carbon content (as well as an ash and catalystcontent) and may be removed periodically from hydromethanation reactor(200) via a char withdrawal line (210) for catalyst recovery and furtherprocessing.

All or a portion of hydromethanation reactor char product stream (54)(typically all of such stream) is processed in a catalyst recovery unit(300) to recover entrained hydromethanation catalyst, and optionallyother value-added by-products such as vanadium and nickel (depending onthe content of the non-gaseous carbonaceous material (10)), to generatea washed char product stream (60) and a recovered hydromethanationcatalyst stream (57).

All or a portion of washed char product stream (60) is processed asdescribed below. The remainder of the washed char, (59), may beprocessed in a boiler to generate steam and power.

In hydromethanation reactor (200), the methane-enriched raw product gastypically passes through the disengagement zone (204) above thefluidized-bed section (202) prior to withdrawal from hydromethanationreactor (200). The disengagement zone (204) may optionally contain, forexample, one or more internal cyclones and/or other entrained particledisengagement mechanisms (not shown). The withdrawn methane-enriched rawproduct gas stream (50) typically comprises at least methane, carbonmonoxide, carbon dioxide, hydrogen, hydrogen sulfide, steam, andentrained solids.

The methane-enriched raw product gas stream (50) is initially treated toremove a substantial portion of the entrained solids, typically via acyclone assembly (for example, one or more internal and/or externalcyclones), which may be followed if necessary by optional additionaltreatments such as venturi scrubbers. In the embodiment as shown in FIG.1, the cyclone assembly comprises an external primary cyclone (360)followed by an external secondary cyclone (370), but other arrangementswould be suitable as well. For example, the cyclone assembly couldcomprise an internal primary cyclone followed by an external secondarycyclone.

The withdrawn methane-enriched raw product gas stream (50), therefore,is to be considered the raw product prior to fines separation,regardless of whether the fines separation takes place internal toand/or external of hydromethanation reactor (200).

As specifically depicted in FIG. 1, the methane-enriched raw product gasstream (50) is passed from hydromethanation reactor (200) to an externalprimary cyclone (360) for separation of the predominant portion ofentrained solids. While primary cyclone (360) is shown as a singleexternal cyclone for simplicity, as indicated above cyclone assembly(360) may be an internal and/or external cyclone and may also be aseries of multiple internal and/or external cyclones.

As shown in FIG. 1, the methane-enriched raw product gas stream (50) istreated in primary cyclone (360) to generate a solids-depletedmethane-enriched raw product gas stream (64) and a recovered primarysolids stream (362).

Recovered primary solids stream (362) is fed back into hydromethanationreactor (200), for example, into one or more portions of fluidized bed(202) via fines recycle line (364). For example, as disclosed inpreviously incorporated US2012/0060417A1, recovered fines are fed backinto lower portion (202 a) of fluidized bed (202) via fines recycle line(364).

The solids-depleted methane-enriched raw product gas stream (64)typically comprises at least methane, carbon monoxide, carbon dioxide,hydrogen, hydrogen sulfide, steam, and ammonia, as well as small amountsof contaminants such as remaining residual entrained fines, and othervolatilized and/or carried material that may be present in thecarbonaceous feedstock. There are typically virtually no (totaltypically less than about 50 ppm) condensable (at ambient conditions)hydrocarbons present in solids-depleted methane-enriched raw product gasstream (64).

Typically, as shown in FIG. 1, the solids-depleted methane-enriched rawproduct gas stream (64) will be fed to a secondary cyclone (370) toremove a substantial portion of any remaining fines, generating afines-cleaned methane-enriched raw product gas stream (70) and arecovered secondary fines stream (66). Recovered secondary fines stream(66) will typically be recycled back to catalyst recovery unit (300).

All or a portion of a recovered secondary fines stream (66) may beco-processed with the withdrawn hydromethanation reactor char productstream (54) in the catalyst recovery unit (300).

The catalyst recovery unit (300) recovers the water-soluble catalyst byconventional solids leaching or washing technologies, extracting aportion of the entrained catalyst to generate a washed product charstream (60) and a recovered hydromethanation catalyst stream (57). Unit(300) may include countercurrent mixer settlers or filter presses withwash zones or any combination of similar solid washing/leaching anddewatering devices. In particular, the catalyst recovery unit (300) maycomprise a quench tank and a quench medium, the treatment comprising thesteps of: quenching the hydromethanation reactor char product stream(54) with the quench medium to extract a portion of the entrainedcatalyst to generate a catalyst-depleted char and liberatedhydromethanation catalyst; withdrawing a stream of catalyst-depletedchar from the catalyst recovery unit (300) as the washed char productstream (60); and withdrawing a stream of liberated hydromethanationcatalyst from the catalyst recovery unit (300) as the recoveredhydromethanation catalyst stream (57).

The hydromethanation catalyst (31) will typically comprise at least aportion of the recovered hydromethanation catalyst stream (57) and amake-up catalyst from a make-up catalyst stream (56).

The fines-cleaned methane-enriched raw product gas stream (70) can betreated in one or more downstream processing steps to recover heatenergy, decontaminate and convert, to one or more value-added productssuch as, for example, substitute natural gas (pipeline quality),hydrogen, carbon monoxide, syngas, ammonia, methanol and othersyngas-derived products, electrical power and steam.

EMBODIMENTS OF THE INVENTION

FIGS. 2 and 3 represent two of the alternative embodiments of thecatalytic gasification process that aim to maximize the overall carbonconversion and power generated from the process.

The washed char stream (60) is further processed in the oxidationreactor (380) with either oxygen or air to convert the residual carbonpresent in the washed char stream (60). The alternative embodimentsdiffer in three ways: (1) the pressure of operation of the oxidationreactor (380), (2) the type of oxidant chosen to convert the carbon, and(3) the molecular ratio of gaseous oxygen to carbon in the feed streamsto the oxidation reactor (380).

The oxidant is an oxygen-containing gas stream, comprising compressedair (87) or oxygen (15 a) from air separation unit (14), and,optionally, diluent steam (12 a). The pressure of the oxidant is greaterthan the operating pressure of the oxidation reactor (380).

The composition of the oxidation reactor product gas (88) depends on themolecular ratio of oxygen to carbon in the streams entering oxidationreactor (380). When the molecular ratio of oxygen to carbon iscontrolled to be less than 0.5 and a sufficient quantity of steam isalso fed to the oxidation reactor (380), the gas (88) is a syngascomprised of hydrogen, carbon monoxide, carbon dioxide, nitrogen andunreacted steam. The mode of operation is known as the partial oxidationmode wherein oxygen is completely consumed within the oxidation reactor(380). On the other hand, when the molecular ratio of oxygen to carbonis controlled to be greater than one (combustion-mode), the gas (88) isa flue gas comprised of carbon dioxide, steam, nitrogen and unreactedoxygen.

First Embodiment of the Invention

FIG. 2 illustrates one of the embodiments of the process for convertinga non-gaseous carbonaceous material ultimately into one or morevalue-added gaseous products where the oxidation reactor is operated ata pressure greater than 791 kPa (100 psig) in syngas-mode. In accordancewith this embodiment, the washed char stream (60) produced by the basichydromethanation process described above and in FIG. 1 is furtherprocessed in an oxidation reactor (380) with steam (12 a) and eitheroxygen (15 a) or compressed air (87) as the oxidant. Air compressor orblower (990) produces compressed air (87) and, optionally, stream (86a), from air (85). The process yields an oxidation reactor ash productstream (89) and an oxidation reactor product gas (88). The oxidation gasis then utilized to produce electric power and steam as described below.

The oxidation reactor ash product stream (89) exiting the oxidationreactor is cooled by a coolant stream (94) to produce a cooled productash (95) and a warmed coolant stream (96). The ash cooler (385) can beany number of direct or indirect heat transfer devices, such as anair-blown fluidized bed, a fluidized bed with steam or water coolingcoils, a direct water quench vessel or a water cooled mechanical screw.

Since the oxidation reactor (380) is operated at high pressure in thisembodiment, it is advantageous to recover and convert the potentialenergy of the pressurized oxidation gases to mechanical work andelectric power. The oxidation reactor product gas (88) produced is richin hydrogen and carbon monoxide and is burned directly in ahigh-pressure gas turbine combustion chamber (992).

Prior to combustion, gas (88) may be processed to remove particulatematter entrained from the oxidation reactor. Suitable mechanisms forremoving particulates from the hot oxidation gas include stage cyclones,candle filters and other similar devices (not shown).

Gas (88) is then reacted with atmospheric air (90) compressed to matchthe pressure of the combustion chamber (992). The air compressor (991)is typically coupled to a gas expander (993). The work recovered by thegas expander (993) drives both the air compressor (991) and an electricgenerator “G” (99 a).

The expander exhaust gas (97) is directed to a heat recovery steamgenerator (975). The heat recovery steam generator (975) consists ofseveral coils to preheat boiler feed water and produce steam, some ofwhich may be extracted for process use and the balance exported as steam(92). The balance of the steam is directed to a steam turbine (980) torecover additional work and generate more electricity by the electricgenerator “G” (99 b), which is coupled to the steam turbine. The steamturbine can be either back pressure or condensing, with multiple stagesand extraction ports if needed. A condensing turbine will maximize theelectric power generated by the steam turbine. The heat recovery steamgenerator (975) also produces flue gas (93).

Second Embodiment of the Invention

FIG. 3 illustrates a second preferred embodiment of the process forconverting a non-gaseous carbonaceous material ultimately into one ormore value-added gaseous products. The oxidant used is compressed air(87) produced by air compressor or blower (990) from air (85); stream(86 a) is also produced. The oxidation reactor (380) may be operated ata high pressure greater than 791 kPa (100 psig) or at a low pressureless than 400 kPa (44 psig). The ratio of air flow rate to the flow rateof the washed char stream (60) produced by the basic hydromethanationprocess described above and in FIG. 1 may be controlled to allowoperation in the syngas-mode or combustion-mode, as described above. Inthe syngas-mode, the oxidation reactor (380) may also be supplied withsteam (12 a) in addition to compressed air (87) to yield an oxidationreactor ash product stream (89) and an oxidation reactor product gas(88). The oxidation reactor ash product stream (89) is cooled in ashcooler (385) by a coolant stream (94) to produce a cooled product ash(95) and a warmed coolant stream (96). The oxidation gas is thenutilized to produce electric power and steam as described below.

When the oxidation reactor (380) is operated at high pressure, thepotential energy of the pressurized oxidation reactor product gas (88)may be converted to mechanical work and electric power in a gas expander(993). The work recovered by the gas expander (993) in this case drivesthe electric generator “G” (99 a).

When the oxidation reactor (380) is operated in syngas-mode, the heatrecovery steam generator (975) will be equipped with an auxiliary burner(not shown) to combust the expander exhaust gas (97) with air (91) toform oxidation products comprised of carbon dioxide and water to allowdischarge to the atmosphere as flue gas (93).

The expander exhaust gas (97) is subsequently processed in the heatrecovery steam generator (975) and steam turbine (980) to produce steam(92) and power (99 b) as described in the first embodiment of theprocess.

In all cases, the flue gas (93) is treated to remove particulate matterand sulfur oxides using standard air pollution control equipment (notshown).

Unit Operations and Processing Details

Hydromethanation (200)

As illustrated in FIG. 1, catalyzed carbonaceous feedstock (31+32), asteam stream (12), and an oxygen stream (15) are introduced intohydromethanation reactor (200).

Char by-product removal from hydromethanation reactor (200) can be atany desired place or places, for example, at the top of fluidized bed(202), at any place within upper portion (202 b) and/or lower portion(202 a) of fluidized bed (202), and/or at or just below hydromethanationreactor fluidizing gas distributor plate (208). As indicated above, thelocation where catalyzed carbonaceous feedstock (31+32) is introducedwill have an influence on the location of a char withdrawal point.

Typically, there will be at least one char withdrawal point at or belowhydromethanation reactor fluidizing gas distributor plate (208) towithdraw char comprising larger or agglomerated particles.

Hydromethanation reactor (200) is typically operated at moderately highpressures and temperatures, requiring introduction of solid streams(e.g., catalyzed carbonaceous feedstock (31+32) and, if present, recyclefines) to the reaction chamber of the reactor while maintaining therequired temperature, pressure and flow rate of the streams. Thoseskilled in the art are familiar with feed inlets to supply solids intothe reaction chambers having high pressure and/or temperatureenvironments, including star feeders, screw feeders, rotary pistons andlock-hoppers. It should be understood that the feed inlets can includetwo or more pressure-balanced elements, such as lock hoppers, whichwould be used alternately. In some instances, the carbonaceous feedstockcan be prepared at pressure conditions above the operating pressure ofthe reactor and, hence, the particulate composition can be directlypassed into the reactor without further pressurization. Gas forpressurization can be an inert gas such as nitrogen, or more typically astream of carbon dioxide that can, for example be recycled from a carbondioxide stream generated by an acid gas removal unit.

Hydromethanation reactor (200) is desirably operated at a moderatetemperature (as compared to “conventional” oxidation-based gasificationprocesses), with an operating temperature of from about 400° F. (about205° C.), or from about 1000° F. (about 538° C.), or from about 1100° F.(about 593° C.), or from about 1200° F. (about 649° C.), to about 1500°F. (about 816° C.), or to about 1400° F. (about 760° C.), or to about1375° F. (about 746° C.); and a pressure of at least about 250 psig(about 1825 kPa, absolute), or at least about 400 psig (about 2860 kPa),or at least about 450 psig (about 3204 kPa). Typically, the pressure canrange up to the levels of mechanical feasibility, for example, up toabout 1200 psig (about 8375 kPa), up to about 1000 psig (about 6996kPa), or to about 800 psig (about 5617 kPa), or to about 700 psig (about4928 kPa), or to about 600 psig (about 4238 kPa), or to about 500 psig(about 3549 kPa). In one embodiment, hydromethanation reactor (200) isoperated at a pressure (first operating pressure) of up to about 600psig (about 4238 kPa), or up to about 550 psig (about 3894 kPa).

Typical gas flow velocities in hydromethanation reactor (200) are fromabout 0.5 ft/sec (about 0.15 m/sec), or from about 1 ft/sec (about 0.3m/sec), to about 2.0 ft/sec (about 0.6 m/sec), or to about 1.5 ft/sec(about 0.45 m/sec).

As oxygen stream (15) is fed into hydromethanation reactor (200), aportion of the carbonaceous feedstock (desirably carbon from thepartially reacted feedstock, by-product char and recycled fines) will beconsumed in an oxidation/combustion reaction, generating heat energy aswell as typically some amounts carbon monoxide and hydrogen (andtypically other gases such as carbon dioxide and steam). The variationof the amount of oxygen supplied to hydromethanation reactor (200)provides an advantageous process control to ultimately maintain thesyngas and heat balance. Increasing the amount of oxygen will increasethe oxidation/combustion, and therefore increase in situ heatgeneration. Decreasing the amount of oxygen will conversely decrease thein situ heat generation. The amount of syngas generated will ultimatelydepend on the amount of oxygen utilized, and higher amounts of oxygenmay result in a more complete combustion/oxidation to carbon dioxide andwater, as opposed to a more partial combustion to carbon monoxide andhydrogen.

The amount of oxygen supplied to hydromethanation reactor (200) must besufficient to combust/oxidize enough of the carbonaceous feedstock togenerate enough heat energy and syngas to meet the heat and syngasdemands of the steady-state hydromethanation reaction.

A portion of the oxygen (second oxygen stream 15 a) is fed into anoxidation reactor (380) along with a portion of the steam (steam stream12 a) to react with by-product char as discussed in further detailbelow.

In one embodiment, the total amount of molecular oxygen that is providedto the hydromethanation reactor (200), or the hydromethanation reactor(200) and oxidation reactor (380) combined, can range from about 0.10,or from about 0.20, or from about 0.25, to about 0.6, or to about 0.5,or to about 0.4, or to about 0.35 weight units (for example, pound orkg) of O₂ per weight unit (for example, pound or kg) of carbonaceousfeedstock (32).

The hydromethanation and oxidation/combustion reactions withinhydromethanation reactor (200) will occur contemporaneously. Dependingon the configuration of hydromethanation reactor (200), the two stepswill typically predominate in separate zones—the hydromethanation inupper portion (202 b) of fluidized bed (202), and theoxidation/combustion in lower portion (202 a) of fluidized bed (202).

Oxygen stream (15) is typically mixed with steam stream (12) and themixture introduced at or near the bottom of fluidized bed (202) in lowerportion (202 a) through hydromethanation reactor fluidizing gasdistributor plate (208) to avoid formation of hot spots in the reactor,to avoid (minimize) combustion of the desired gaseous products generatedwithin hydromethanation reactor (200), and to improve the safety of theoperation. Feeding the catalyzed carbonaceous feedstock (31+32) with anelevated moisture content, and particularly into lower portion (202 a)of fluidized bed (202), also assists in heat dissipation and theavoidance of formation of hot spots in reactor (200), as disclosed inpreviously incorporated US2012/0102837A1.

Oxygen stream (15) can be fed into hydromethanation reactor (200) by anysuitable means such as direct injection of purified oxygen, oxygen-airmixtures, oxygen-steam mixtures, or oxygen-inert gas mixtures into thereactor. See, for instance, U.S. Pat. No. 4,315,753 and Chiaramonte etal., Hydrocarbon Processing, September 1982, pp. 255-257. Oxygen stream(15) and second oxygen stream (15 a), which splits off from oxygenstream (15), is supplied by an air separation unit (14).

Oxygen stream (15) is typically generated via standard air-separationtechnologies, and will be fed mixed with steam, and introduced at atemperature above about 250° F. (about 121° C.), to about 400° F. (about204° C.), or to about 350° F. (about 177° C.), or to about 300° F.(about 149° C.), and at a pressure at least slightly higher than presentin hydromethanation reactor (200). The steam in oxygen stream (15)should be non-condensable during transport of oxygen stream (15) tohydromethanation reactor (200), so oxygen stream (15) may need to betransported at a lower pressure then pressurized (compressed) just priorto introduction into hydromethanation reactor (200).

As indicated above, the hydromethanation reaction has a steam demand, aheat demand and a syngas demand. These conditions in combination areimportant factors in determining the operating conditions for thehydromethanation reaction as well as the remainder of the process.

For example, the hydromethanation reaction requires a theoretical molarratio of steam to carbon (in the feedstock) of at least about 1.Typically, however, the molar ratio is greater than about 1, or fromabout 1.5 (or greater), to about 6 (or less), or to about 5 (or less),or to about 4 (or less), or to about 3 (or less), or to about 2 (orless). The moisture content of the catalyzed carbonaceous feedstock(31+32), moisture generated from the carbonaceous feedstock in thehydromethanation reactor (200), and steam included in the steam stream(12), oxygen stream (15), oxidation reactor product gas stream (88) andrecycle fines stream(s) all contribute steam for the hydromethanationreaction. The steam in steam stream (12), oxygen stream (15) andoxidation gas stream (88) should be sufficient to at least substantiallysatisfy (or at least satisfy) the “steam demand” of the hydromethanationreaction.

As also indicated above, the hydromethanation reaction is essentiallythermally balanced but, due to process heat losses and other energyrequirements (for example, vaporization of moisture on the feedstock),some heat must be generated in situ (in hydromethanation reactor (200))to maintain the thermal balance (the heat demand). The partialcombustion/oxidation of carbon in the presence of the oxygen introducedinto hydromethanation reactor (200) from oxygen stream (15) should besufficient to at least substantially satisfy (or at least satisfy) boththe heat and syngas demand of the hydromethanation reaction.

The gas utilized in hydromethanation reactor (200) for pressurizationand reaction of the catalyzed carbonaceous feedstock (31+32) comprisesthe steam stream (12), oxygen stream (15) and, optionally, additionalnitrogen, air, or inert gases such as argon, which can be supplied tohydromethanation reactor (200) according to methods known to thoseskilled in the art. As a consequence, steam stream (12) and oxygenstream (15) must be provided at a higher pressure which allows them toenter hydromethanation reactor (200).

Steam stream (12) can be at a temperature as low as the saturation pointat the feed pressure, but it is desirable to feed at a temperature abovethis to avoid the possibility of any condensation occurring. Typicalfeed temperatures of superheated steam stream (12) are from about 500°F. (about 260° C.), or from about 600° F. (about 316° C.), or from about700° F. (about 371° C.), to about 950° F. (about 510° C.), or to about900° F. (about 482° C.). Typical feed pressures of steam stream (12) areabout 25 psi (about 172 kPa) or greater than the pressure withinhydromethanation reactor (200).

The actual temperature and pressure of steam stream (12) will ultimatelydepend on the level of heat recovery from the process and the operatingpressure within hydromethanation reactor (200), as discussed below.

When steam stream (12) and oxygen stream (15) are combined for feedinginto lower section (202 a) of fluidized bed (202), the temperature ofthe combined stream will be controlled by the temperature of steamstream (12), and will typically range from about from about from about400° F. (about 204° C.), or from about 450° F. (about 232° C.), to about800° F. (about 455° C.), or to about 600° F. (about 316° C.).

The temperature in hydromethanation reactor (200) can be controlled, forexample, by controlling the amount and temperature of steam stream (12)as well as the amount of oxygen supplied to hydromethanation reactor(200).

In steady-state operation, steam for hydromethanation reactor (200) andoxidation reactor (380) is desirably solely generated from other processoperations through process heat capture (such as generated in a wasteheat boiler, generally referred to as “process steam” or“process-generated steam”, and referenced as steam source (11)),specifically from the cooling of the raw product gas in a heat exchangerunit. Additional steam can be generated for other portions of theoverall process, such as disclosed, for example, in previouslyincorporated US2010/0287835A1 and US2012/0046510A1, and as shown inFIGS. 2 and 3 discussed below.

The result of the hydromethanation reaction is a methane-enriched rawproduct, which is withdrawn from hydromethanation reactor (200) asmethane-enriched raw product gas stream (50) typically comprising CH₄,CO₂, H₂, CO, H₂S, unreacted steam and, optionally, other contaminantssuch as entrained solids, NH₃, COS, HCN and/or elemental mercury vapor,depending on the nature of the carbonaceous material utilized forhydromethanation.

The non-gaseous carbonaceous materials (10) useful in these processesinclude, for example, a wide variety of biomass and non-biomassmaterials. The carbonaceous feedstock (32) is derived from one or morenon-gaseous carbonaceous materials (10), which are processed in afeedstock preparation unit (100) as discussed below.

The hydromethanation catalyst (31) can comprise one or more catalystspecies, as discussed below.

The carbonaceous feedstock (32) and the hydromethanation catalyst (31)are typically intimately mixed (i.e., to provide a catalyzedcarbonaceous feedstock (31+32)) before provision to the hydromethanationreactor (200), but they can be fed separately as well.

The hot gas effluent leaving the reaction chamber of thehydromethanation reactor (200) can pass through a fines remover unit(such as cyclone assembly (360)), incorporated into and/or external ofthe hydromethanation reactor (200), which serves as a disengagementzone. Particles too heavy to be entrained by the gas leaving thehydromethanation reactor (200) (i.e., fines) are returned to thehydromethanation reactor (200), for example, to the reaction chamber(e.g., fluidized bed (202)).

Residual entrained solids are substantially removed by any suitabledevice such as internal and/or external cyclone separators to generate asolids-depleted methane-enriched raw product gas stream (64). Asdiscussed above, at least a portion of these fines can be returned tofluidized bed (202) via recycle line (364). Any remaining recoveredfines can be processed to recover alkali metal catalyst, and/or combinedat some stage with carbonaceous feedstock (32), and/or directly recycledback to feedstock preparation as described in previously incorporatedUS2009/0217589A1.

Removal of a “substantial portion” of fines means that an amount offines is removed from the resulting gas stream such that downstreamprocessing is not adversely affected; thus, at least a substantialportion of fines should be removed. Some minor level of ultrafinematerial may remain in the resulting gas stream to the extent thatdownstream processing is not significantly adversely affected.Typically, at least about 90 wt %, or at least about 95 wt %, or atleast about 98 wt %, of the fines of a particle size greater than about20 μm, or greater than about 10 μm, or greater than about 5 μm, areremoved.

Additional residual entrained fines may be removed from thesolids-depleted methane-enriched raw product gas stream (64) by anysuitable device such as internal and/or external cyclone separators suchas external secondary cyclone (370), optionally followed by scrubbers.The resulting fines-cleaned methane-enriched raw product gas stream (70)can be further processed for heat recovery and/orpurification/conversion as required to achieve a desired end product, asdisclosed in the numerous previously incorporated disclosures set forthabove in the “Hydromethanation” section. Reference may be had to thosedisclosures for further details.

Oxidation Reactor (380)

Oxidation reactor (380) can be any type of oxidation/combustion reactorsuitable for reacting the by-product char with oxygen or air and steamunder the conditions specified by the embodiments of this invention.Suitable reactors are in a general sense well-known to those of ordinaryskill in the relevant art.

Preferably, the oxidation reactor (380) is a fluidized-bed reactorwhich, in operation, contains a fluidized bed of the by-product charfluidized by steam (12 a) and oxygen (15 a) or compressed air (87) fedinto the reactor through oxidation reactor fluidizing gas distributorplate (230).

The oxidation reactor may include internal and external cyclones (notshown) to remove particles from the oxidation gas. The recoveredparticles maybe recycled to oxidation reactor as is typical for a fluidbed or the particles may be directed to the char cooling system fordisposal.

Oxidation reactor (380) operates at a temperature of about 1800° F.(about 982° C.) or less, or about 1450° F. (about 788° C.) or less, orabout 1400° F. (about 760° C.) or less. The pressure of operation of theoxidation reactor (380) may be chosen independent of thehydromethanation reactor (200) since they are decoupled. In theoxidation reactor (380), all or a portion of the carbon content of theby-product char is reacted with steam and oxygen or air to produce acarbon-depleted char and an oxidation gas comprising steam, carbonmonoxide, carbon dioxide, hydrogen, and nitrogen (if air is used as anoxidant). A stream of the oxidation gas is withdrawn from oxidationreactor (380) as oxidation reactor product gas stream (88), and a streamof carbon-depleted char is withdrawn from oxidation reactor as oxidationreactor ash product stream (89).

The oxidation reactor product gas stream (88) is a “medium BTU gas” ifthe constituents are primarily steam, CO, CO₂ and H₂. Gas stream (88) isa “low BTU gas” if the constituents are primarily nitrogen, CO, CO₂, H₂and steam. Gas stream (88) is a “flue gas” with no chemical energycontent if the constituents are primarily nitrogen, steam and CO₂.

The oxidation reactor product gas (88) is utilized to produce steam andelectric power, which are the main objectives of this invention and thedetails of which are described elsewhere.

Air Compressor (990)

An air compressor or blower (990) compresses air (85), supplyingcompressed air (87) to the oxidation reactor (380). A number ofcommercially available machines are available for this application.Compressors are staged to achieve progressively higher pressures.Centrifugal compressors or turboblowers are widely used to handle largevolumes of gas from 0.5 up to several hundred psi. For air, the typicalpressure ratio per stage is 1.4. The gas is cooled between stages as thework of compression will cause the air to increase in temperature.

Rotary compressors and blowers, screw compressors and reciprocatingcompressors can also be used for this application. The choice of machineis application dependent; important selection factors are the requiredpressure ratio, volume of air and machine costs and efficiency.

When the pressure of the compressed air (87) exceeds the pressure of theair required for the air separation unit (14), all or a portion of thecompressed air required for air separation may be extracted as stream(86 a). Similarly, when the air compressor (991) is operated at apressure higher than that of the air required for the air separationunit (14), all or a portion of the compressed air required for airseparation may be extracted as stream (86 b).

Ash Cooler (385)

The ash cooler is a device to reduce the temperature and often thepressure of the solid product from the oxidation reactor (380) so thematerial can be safely handled and disposed. A number of differentdevices can be employed to achieve this goal. The simplest device is adirect quench system whereby the char or ash is dropped into a pool ofcooler water. The water drops the temperature of the char by directcontact and the resultant water/solids slurry can then be depressurizedand further cooled if need be. The slurry can be depressurized byexpanding through an abrasion resistant control valve or let downthrough an arrangement of lock hoppers.

Another means of cooling the char or ash is to direct the material intoa fluid bed with internal cooling coils. The bed may be fluidized withcold air and the coils may contain cooling fluid such as water. The heattransfer between coils and solids in a fluid bed is known to occur withhigh efficiency. The main advantage of this method is the char or ashcan be withdrawn dry.

A third option is to cool the char by a cooling screw. This deviceemploys a transfer screw encased in a water cooled jacket. The screw mayalso be internally cooled. Again, the char or ash can be kept dry as itis cooled If the char is at pressure a separate method to depressurizethe cooled char may be required, such as lock hoppers or slide valves.

Catalyst Recovery (300)

The hydromethanation reactor char product stream (54) typicallycomprises inorganic ash and entrained catalyst, as well as residualcarbon. This char is passed to a catalyst recovery unit (300).

In certain embodiments, when the hydromethanation catalyst is an alkalimetal, the alkali metal in the hydromethanation reactor char productstream (54) can be recovered to produce a recovered hydromethanationcatalyst stream (57), and any unrecovered catalyst can be compensated bya make-up catalyst stream (56) (see, for example, previouslyincorporated US2009/0165384A1). The more alumina plus silica that is inthe feedstock, the more costly it is to obtain a higher alkali metalrecovery.

Ultimately, the recovered hydromethanation catalyst stream (57) can bedirected to the catalyst application unit (350) for reuse of the alkalimetal catalyst.

Other particularly useful recovery and recycling processes are describedin U.S. Pat. No. 4,459,138, as well as previously incorporatedUS2007/0277437A1 US2009/0165383A1, US2009/0165382A1, US2009/0169449A1and US2009/0169448A1. Reference can be had to those documents forfurther process details.

The recycle of catalyst can be to one or a combination of catalystloading processes. For example, all of the recycled catalyst can besupplied to one catalyst loading process, while another process utilizesonly makeup catalyst. The levels of recycled versus makeup catalyst canalso be controlled on an individual basis among catalyst loadingprocesses.

The washed char product stream (60) can also be treated for recovery ofother by-products, such as vanadium and/or nickel, in addition tocatalyst recovery, as disclosed in previously incorporatedUS2011/0262323A1 and US2012/0213680A1.

As indicated above, recovered secondary fines stream (66) can beco-treated in catalyst recovery unit (300).

In this embodiment, the washed char stream (60) is further processed soit can be fed into the oxidation reactor (380). These processing stepscould include direct thermal drying, pelletization and thermal drying orbriquetting.

Preparation of Carbonaceous Feedstocks

Carbonaceous Materials Processing (100)

Particulate carbonaceous materials, such as biomass and non-biomass, canbe prepared via crushing and/or grinding, either separately or together,according to any methods known in the art, such as impact crushing andwet or dry grinding to yield one or more carbonaceous particulates.Depending on the method utilized for crushing and/or grinding of thecarbonaceous material sources, the resulting carbonaceous particulatesmay be sized (i.e., separated according to size) to provide thecarbonaceous feedstock (32) for use in catalyst application unit (350)to form a catalyzed carbonaceous feedstock (31+32) for thehydromethanation reactor (200).

Any method known to those skilled in the art can be used to size theparticulates. For example, sizing can be performed by screening orpassing the particulates through a screen or number of screens.Screening equipment can include grizzlies, bar screens, and wire meshscreens. Screens can be static or incorporate mechanisms to shake orvibrate the screen. Alternatively, classification can be used toseparate the carbonaceous particulates. Classification equipment caninclude ore sorters, gas cyclones, hydrocyclones, rake classifiers,rotating trommels or fluidized classifiers. The carbonaceous materialscan be also sized or classified prior to grinding and/or crushing.

The carbonaceous particulate can be supplied as a fine particulatehaving an average particle size of from about 25 microns, or from about45 microns, up to about 2500 microns, or up to about 500 microns. Oneskilled in the art can readily determine the appropriate particle sizefor the carbonaceous particulates. For example, for a fluidized bedreactor, such carbonaceous particulates can have an average particlesize which enables incipient fluidization of the carbonaceous materialsat the gas velocity used in the fluidized bed reactor. Desirableparticle size ranges for the hydromethanation reactor (200) are in theGeldart A and Geldart B ranges (including overlap between the two),depending on fluidization conditions, typically with limited amounts offine (below about 25 microns) and coarse (greater than about 250microns) material.

Additionally, certain carbonaceous materials, for example, corn stoverand switchgrass, and industrial wastes, such as saw dust, either may notbe amenable to crushing or grinding operations, or may not be suitablefor use as such, for example due to ultrafine particle sizes. Suchmaterials may be formed into pellets or briquettes of a suitable sizefor crushing or for direct use in, for example, a fluidized bed reactor.Generally, pellets can be prepared by compaction of one or morecarbonaceous material; see for example, previously incorporatedUS2009/0218424A1. In other examples, a biomass material and a coal canbe formed into briquettes as described in U.S. Pat. Nos. 4,249,471,4,152,119 and 4,225,457. Such pellets or briquettes can be usedinterchangeably with the preceding carbonaceous particulates in thefollowing discussions.

Additional feedstock processing steps may be necessary depending on thequalities of carbonaceous material sources. Biomass may contain highmoisture contents, such as green plants and grasses, and may requiredrying prior to crushing. Municipal wastes and sewages also may containhigh moisture contents which may be reduced, for example, by use of apress or roll mill (e.g., U.S. Pat. No. 4,436,028). Likewise,non-biomass, such as high-moisture coal, can require drying prior tocrushing. Some caking coals can require partial oxidation to simplifyoperation. Non-biomass feedstocks deficient in ion-exchange sites, suchas anthracites or petroleum cokes, can be pre-treated to createadditional ion-exchange sites to facilitate catalyst loading and/orassociation. Such pre-treatments can be accomplished by any method knownto the art that creates ion-exchange capable sites and/or enhances theporosity of the feedstock (see, for example, previously incorporatedU.S. Pat. No. 4,468,231 and GB1599932). Oxidative pre-treatment can beaccomplished using any oxidant known to the art.

The ratio and types of the carbonaceous materials in the carbonaceousparticulates can be selected based on technical considerations,processing economics, availability, and proximity of the non-biomass andbiomass sources. The availability and proximity of the sources for thecarbonaceous materials can affect the price of the feeds, and thus theoverall production costs of the catalytic gasification process. Forexample, the biomass and the non-biomass materials can be blended in atabout 5:95, about 10:90, about 15:85, about 20:80, about 25:75, about30:70, about 35:65, about 40:60, about 45:55, about 50:50, about 55:45,about 60:40, about 65:35, about 70:20, about 75:25, about 80:20, about85:15, about 90:10, or about 95:5 by weight on a wet or dry basis,depending on the processing conditions.

Significantly, the carbonaceous material sources, as well as the ratioof the individual components of the carbonaceous particulates, forexample, a biomass particulate and a non-biomass particulate, can beused to control other material characteristics of the carbonaceousparticulates. Non-biomass materials, such as coals, and certain biomassmaterials, such as rice hulls, typically include significant quantitiesof inorganic matter including calcium, alumina and silica which forminorganic oxides (i.e., ash) in the catalytic gasifier. At temperaturesabove about 500° C. to about 600° C., potassium and other alkali metalscan react with the alumina and silica in ash to form insoluble alkalialuminosilicates. In this form, the alkali metal is substantiallywater-insoluble and inactive as a catalyst. To prevent buildup of theresidue in the hydromethanation reactor (200), as described above asolid purge of hydromethanation reactor char product stream (54)comprising ash, unreacted carbonaceous material, and various othercompounds (such as alkali metal compounds, both water soluble and waterinsoluble) is withdrawn and processed.

In preparing the carbonaceous particulates, the ash content of thevarious carbonaceous materials can be selected to be, for example, about20 wt % or less, or about 15 wt % or less, or about 10 wt % or less, orabout 5 wt % or less, depending on, for example, the ratio of thevarious carbonaceous materials and/or the starting ash in the variouscarbonaceous materials. In other embodiments, the resulting thecarbonaceous particulates can comprise an ash content ranging from about5 wt %, or from about 10 wt %, to about 20 wt %, or to about 15 wt %,based on the weight of the carbonaceous particulate. In otherembodiments, the ash content of the carbonaceous particulate cancomprise less than about 20 wt %, or less than about 15 wt %, or lessthan about 10 wt %, or less than about 8 wt %, or less than about 6 wt %alumina, based on the weight of the ash. In certain embodiments, thecarbonaceous particulates can comprise an ash content of less than about20 wt %, based on the weight of processed feedstock where the ashcontent of the carbonaceous particulate comprises less than about 20 wt% alumina, or less than about 15 wt % alumina, based on the weight ofthe ash.

Such lower alumina values in the carbonaceous particulates allow for,ultimately, decreased losses of catalysts, and particularly alkali metalcatalysts, in the hydromethanation portion of the process. As indicatedabove, alumina can react with alkali source to yield an insoluble charcomponent comprising, for example, an alkali aluminate oraluminosilicate. Such insoluble char component can lead to decreasedcatalyst recovery (i.e., increased catalyst loss), and thus, requireadditional costs of make-up catalyst in the overall process.

Additionally, the resulting carbonaceous particulates can have asignificantly higher % carbon, and thus btu/lb value and methane productper unit weight of the carbonaceous particulate. In certain embodiments,the resulting carbonaceous particulates can have a carbon contentranging from about 75 wt %, or from about 80 wt %, or from about 85 wt%, or from about 90 wt %, up to about 95 wt %, based on the combinedweight of the non-biomass and biomass.

In one example, a non-biomass and/or biomass is wet ground and sized(e.g., to a particle size distribution of from about 25 to about 2500μm) and then drained of its free water (i.e., dewatered) to a wet cakeconsistency. Examples of suitable methods for the wet grinding, sizing,and dewatering are known to those skilled in the art; for example, seepreviously incorporated US2009/0048476A1. The filter cakes of thenon-biomass and/or biomass particulates formed by the wet grinding inaccordance with one embodiment of the present disclosure can have amoisture content ranging from about 40% to about 60%, or from about 40%to about 55%, or below 50%. It will be appreciated by one of ordinaryskill in the art that the moisture content of dewatered wet groundcarbonaceous materials depends on the particular type of carbonaceousmaterials, the particle size distribution, and the particular dewateringequipment used. Such filter cakes can be thermally treated to produceone or more reduced moisture carbonaceous particulates.

Each of the one or more carbonaceous particulates can have a uniquecomposition, as described above. For example, two carbonaceousparticulates can be utilized, where a first carbonaceous particulatecomprises one or more biomass materials and the second carbonaceousparticulate comprises one or more non-biomass materials. Alternatively,a single carbonaceous particulate comprising one or more carbonaceousmaterials utilized.

Catalyst Loading for Hydromethanation (350)

The hydromethanation catalyst is potentially active for catalyzing atleast reactions (I), (II) and (III) described above. Such catalysts arein a general sense well known to those of ordinary skill in the relevantart and may include, for example, alkali metals, alkaline earth metalsand transition metals, and compounds and complexes thereof. Typically,the hydromethanation catalyst comprises at least an alkali metal, suchas disclosed in many of the previously incorporated references.

For the hydromethanation reaction, the one or more carbonaceousparticulates are typically further processed to associate at least onehydromethanation catalyst, typically comprising a source of at least onealkali metal, to generate a catalyzed carbonaceous feedstock (31+32). Ifa liquid carbonaceous material is used, the hydromethanation catalystmay for example be intimately mixed into the liquid carbonaceousmaterial.

The carbonaceous particulate provided for catalyst loading can be eithertreated to form a catalyzed carbonaceous feedstock (31+32) which ispassed to the hydromethanation reactor (200), or split into one or moreprocessing streams, where at least one of the processing streams isassociated with a hydromethanation catalyst to form at least onecatalyst-treated feedstock stream. The remaining processing streams canbe, for example, treated to associate a second component therewith.Additionally, the catalyst-treated feedstock stream can be treated asecond time to associate a second component therewith. The secondcomponent can be, for example, a second hydromethanation catalyst, aco-catalyst, or other additive.

In one example, the primary hydromethanation catalyst (alkali metalcompound) can be provided to the single carbonaceous particulate (e.g.,a potassium and/or sodium source), followed by a separate treatment toprovide one or more co-catalysts and additives (e.g., a calcium source)to the same single carbonaceous particulate to yield the catalyzedcarbonaceous feedstock (31+32). For example, see previously incorporatedUS2009/0217590A1 and US2009/0217586A1.

The hydromethanation catalyst and second component can also be providedas a mixture in a single treatment to the single second carbonaceousparticulate to yield the catalyzed carbonaceous feedstock (31+32).

When one or more carbonaceous particulates are provided for catalystloading, then at least one of the carbonaceous particulates isassociated with a hydromethanation catalyst to form at least onecatalyst-treated feedstock stream. Further, any of the carbonaceousparticulates can be split into one or more processing streams asdetailed above for association of a second or further componenttherewith. The resulting streams can be blended in any combination toprovide the catalyzed carbonaceous feedstock (31+32), provided at leastone catalyst-treated feedstock stream is utilized to form the catalyzedfeedstock stream.

In one embodiment, at least one carbonaceous particulate is associatedwith a hydromethanation catalyst and optionally, a second component. Inanother embodiment, each carbonaceous particulate is associated with ahydromethanation catalyst and optionally, a second component.

Any methods known to those skilled in the art can be used to associateone or more hydromethanation catalysts with any of the carbonaceousparticulates and/or processing streams. Such methods include but are notlimited to, admixing with a solid catalyst source and impregnating thecatalyst onto the processed carbonaceous material. Several impregnationmethods known to those skilled in the art can be employed to incorporatethe hydromethanation catalysts. These methods include but are notlimited to, incipient wetness impregnation, evaporative impregnation,vacuum impregnation, dip impregnation, ion exchanging, and combinationsof these methods.

In one embodiment, an alkali metal hydromethanation catalyst can beimpregnated into one or more of the carbonaceous particulates and/orprocessing streams by slurrying with a solution (e.g., aqueous) of thecatalyst in a loading tank. When slurried with a solution of thecatalyst and/or co-catalyst, the resulting slurry can be dewatered toprovide a catalyst-treated feedstock stream, again typically, as a wetcake. The catalyst solution can be prepared from any catalyst source inthe present processes, including fresh or make-up catalyst and recycledcatalyst or catalyst solution. Methods for dewatering the slurry toprovide a wet cake of the catalyst-treated feedstock stream includefiltration (gravity or vacuum), centrifugation, and a fluid press.

In another embodiment, as disclosed in previously incorporatedUS2010/0168495A1, the carbonaceous particulates are combined with anaqueous catalyst solution to generate a substantially non-draining wetcake, then mixed under elevated temperature conditions and finally driedto an appropriate moisture level.

One particular method suitable for combining a coal particulate and/or aprocessing stream comprising coal with a hydromethanation catalyst toprovide a catalyst-treated feedstock stream is via ion exchange asdescribed in previously incorporated US2009/0048476A1 andUS2010/0168494A1. Catalyst loading by ion exchange mechanism can bemaximized based on adsorption isotherms specifically developed for thecoal, as discussed in the incorporated reference. Such loading providesa catalyst-treated feedstock stream as a wet cake. Additional catalystretained on the ion-exchanged particulate wet cake, including inside thepores, can be controlled so that the total catalyst target value can beobtained in a controlled manner. The total amount of catalyst loaded canbe controlled by controlling the concentration of catalyst components inthe solution, as well as the contact time, temperature and method, asdisclosed in the aforementioned incorporated references, and as canotherwise be readily determined by those of ordinary skill in therelevant art based on the characteristics of the starting coal.

In another example, one of the carbonaceous particulates and/orprocessing streams can be treated with the hydromethanation catalyst anda second processing stream can be treated with a second component (seepreviously incorporated US2007/0000177A1).

The carbonaceous particulates, processing streams, and/orcatalyst-treated feedstock streams resulting from the preceding can beblended in any combination to provide the catalyzed second carbonaceousfeedstock, provided at least one catalyst-treated feedstock stream isutilized to form the catalyzed carbonaceous feedstock (31+32).Ultimately, the catalyzed carbonaceous feedstock (31+32) is passed ontothe hydromethanation reactor(s) (200).

Generally, each catalyst loading unit comprises at least one loadingtank to contact one or more of the carbonaceous particulates and/orprocessing streams with a solution comprising at least onehydromethanation catalyst, to form one or more catalyst-treatedfeedstock streams. Alternatively, the catalytic component may be blendedas a solid particulate into one or more carbonaceous particulates and/orprocessing streams to form one or more catalyst-treated feedstockstreams.

Typically, when the hydromethanation catalyst is solely or substantiallyan alkali metal, it is present in the catalyzed carbonaceous feedstockin an amount sufficient to provide a ratio of alkali metal atoms tocarbon atoms in the catalyzed carbonaceous feedstock ranging from about0.01, or from about 0.02, or from about 0.03, or from about 0.04, toabout 0.10, or to about 0.08, or to about 0.07, or to about 0.06.

Suitable alkali metals are lithium, sodium, potassium, rubidium, cesium,and mixtures thereof. Particularly useful are potassium sources.Suitable alkali metal compounds include alkali metal carbonates,bicarbonates, formates, oxalates, amides, hydroxides, acetates, orsimilar compounds. For example, the catalyst can comprise one or more ofsodium carbonate, potassium carbonate, rubidium carbonate, lithiumcarbonate, cesium carbonate, sodium hydroxide, potassium hydroxide,rubidium hydroxide or cesium hydroxide, and particularly, potassiumcarbonate and/or potassium hydroxide.

Optional co-catalysts or other catalyst additives may be utilized, suchas those disclosed in the previously incorporated references.

The one or more catalyst-treated feedstock streams that are combined toform the catalyzed carbonaceous feedstock typically comprise greaterthan about 50%, greater than about 70%, or greater than about 85%, orgreater than about 90% of the total amount of the loaded catalystassociated with the catalyzed carbonaceous feedstock (31+32). Thepercentage of total loaded catalyst that is associated with the variouscatalyst-treated feedstock streams can be determined according tomethods known to those skilled in the art. Separate carbonaceousparticulates, catalyst-treated feedstock streams, and processing streamscan be blended appropriately to control, for example, the total catalystloading or other qualities of the catalyzed carbonaceous feedstock(31+32), as discussed previously. The appropriate ratios of the variousstream that are combined will depend on the qualities of thecarbonaceous materials comprising each as well as the desired propertiesof the catalyzed carbonaceous feedstock (31+32). For example, a biomassparticulate stream and a catalyzed non-biomass particulate stream can becombined in such a ratio to yield a catalyzed carbonaceous feedstock(31+32) having a predetermined ash content, as discussed previously.

Any of the preceding catalyst-treated feedstock streams, processingstreams, and processed feedstock streams, as one or more dryparticulates and/or one or more wet cakes, can be combined by anymethods known to those skilled in the art including, but not limited to,kneading, and vertical or horizontal mixers, for example, single or twinscrew, ribbon, or drum mixers. The resulting catalyzed carbonaceousfeedstock (31+32) can be stored for future use or transferred to one ormore feed operations for introduction into the hydromethanationreactor(s). The catalyzed carbonaceous feedstock can be conveyed tostorage or feed operations according to any methods known to thoseskilled in the art, for example, a screw conveyer or pneumatictransport.

In one embodiment, the carbonaceous feedstock as fed to thehydromethanation reactor contains an elevated moisture content of fromgreater than 10 wt %, or about 12 wt % or greater, or about 15 wt % orgreater, to about 25 wt % or less, or to about 20 wt % or less (based onthe total weight of the carbonaceous feedstock), to the extent that thecarbonaceous feedstock is substantially free-flowing (see previouslyincorporated US2012/0102837A1).

The term “substantially free-flowing” as used herein means thecarbonaceous feedstock particulates do not agglomerate under feedconditions due to moisture content. Desirably, the moisture content ofthe carbonaceous feedstock particulates is substantially internallycontained so that there is minimal (or no) surface moisture.

A suitable substantially free-flowing catalyzed carbonaceous feedstock(31+32) can be produced in accordance with the disclosures of previouslyincorporated US2010/0168494A1 and US2010/0168495A1, where the thermaltreatment step (after catalyst application) referred to in thosedisclosures can be minimized (or even potentially eliminated).

To the extent necessary, excess moisture can be removed from thecatalyzed carbonaceous feedstock (31+32). For example, the catalyzedcarbonaceous feedstock (31+32) may be dried with a fluid bed slurrydrier (i.e., treatment with superheated steam to vaporize the liquid),or the solution thermally evaporated or removed under a vacuum, or undera flow of an inert gas, to provide a catalyzed carbonaceous feedstockhaving a the required residual moisture content.

Gas Processing

Heat Exchanger System (400) and High-Pressure Steam Stream (40)

The fines-cleaned methane-enriched raw product gas stream (70) leavingthe hydromethanation reactor (200) contains a very small amount of finesafter being processed by the highly-efficient system of cyclones. FIG. 4illustrates the further processing of the methane-enriched raw productgas stream to make the final product, i.e., pipeline-quality substitutenatural gas (SNG).

The temperature and pressure of stream (70) are dictated by the chosenoperating conditions. The pressure ranges from 250 to 1000 psig (1825 to6996 kPa) and is preferably between 500 and 650 psig (3549 to 4583 kPa).The temperature ranges from 1100 to 1500° F. (593 to 816° C.) and ispreferably between 1250 and 1350° F. (677 to 732° C.).

Referring to FIG. 4, fines-cleaned methane-rich raw product gas stream(70) is routed to a heat exchanger or boiler system (400), optionallycomprising a superheater section (not shown), to recover thermal energyin the form of high-pressure steam stream (40) by vaporization of boilerfeed water (39 a). The pressure of the steam is at least 25 to 50 psig(172 to 345 kPa) higher than the pressure of the hydromethanationreactor (200). Steam stream (40) is preferably superheated to 750 to950° F. (399 to 510° C.) to maximize the thermal efficiency of thehydromethanation reactor (200). In the absence of the superheatersection of heat exchanger system (400) saturated steam may be producedat the pressure stated above for heat exchanger system (400), ifrequired. Steam stream (40) is sent to the steam distribution system(11).

Cooled methane-enriched raw product stream (71) is the methane-enrichedraw product leaving the heat exchanger system (400) that has been cooledto about 550° F. (288° C.). It is further cooled against boiler feedwater (39 b) to 370-400° F. (188 to 204° C.) in intermediate-pressureheat exchanger or boiler (410) to generate intermediate-pressuresaturated steam (41) at 150 psig (1136 kPa), which is close to the dewpoint of syngas under those conditions. Intermediate-pressure saturatedsteam stream (41) is sent to the steam distribution system (11).

Methane-enriched raw product gas stream (72) leavingintermediate-pressure boiler (410) is scrubbed in a hot gas scrubber(420) with recycled process condensate (not shown), which is obtainedfrom an ammonia recovery system (600) (as discussed below), to removeany traces of fine particulate matter that has escaped the cyclones. Ableed stream (not shown) from the hot gas scrubber (420) containing thefine particulate matter is routed to the catalyst recovery system (300)(not shown). A particle-free cooled gas (73) exits the hot gas scrubber(420).

Gas Conversion/Purification

Referring to FIG. 4, raw product conversion/purification will typicallycomprise water-gas shift reactors (700), low temperature gas cooling(450), ammonia recovery (600), acid gas removal (800) and methanation(950).

Water-Gas Shift System (700)

The particle-free cooled gas stream (73) exiting the hot gas scrubber(420) is split into two streams, a shift inlet stream (73 a) and a shiftbypass stream (73 b). Shift inlet stream (73 a) is sent to the water-gasshift system (700) where a portion of the stream is reheated to 450 to550° F. (232 to 288° C.) and then passed over a sour-shift catalyst(typically cobalt-molybdenum) in a fixed bed reactor to convert aportion of the carbon monoxide and steam to hydrogen and carbon dioxide,forming shift outlet stream (73 c). Since the shift reaction isexothermic, the shift outlet stream (73 c) from the shift reactorexchanges heat with shift inlet stream (73 a) to recover energy. Shiftbypass stream (73 b) bypasses water-gas shift system (700) and iscombined with shift outlet stream (73 c) to form hydrogen-enriched rawproduct gas stream (74). The fraction of particle-free cooled gas (73)that is shifted is controlled to maintain a ratio hydrogen to carbonmonoxide of approximately 3:1 in the combined gas stream (74).

Methods and reactors for performing the water-gas shift reaction on aCO-containing gas stream are well known to those of skill in the art. Anexample of a suitable shift reactor is illustrated in U.S. Pat. No.7,074,373, hereby incorporated by reference, although other designsknown to those of skill in the art are also effective.

Low-Temperature Gas Cooling (450)

After leaving the water-gas shift reactor system (700),hydrogen-enriched raw product gas stream (74) is cooled in a series ofheat exchangers within low-temperature gas cooling system (450) tofurther reduce the temperature to 120° F. (49° C.) to produce a dry rawgas stream (75). The low-temperature gas cooling system (450) willtypically comprise first and second knock-out drums, an air cooler, anda trim cooler.

Hydrogen-enriched raw product gas stream (74), initially at about 475°F. (246° C.), is first cooled against boiler feed water (39 c) togenerate medium-pressure steam (42) at 50 psig (446 kPa) andlow-pressure steam (not shown) at two levels: 30 psig (308 kPa) and 15psig (205 kPa). Recovery of low-grade heat allows heat integration withother parts of the process where steam at these pressure levels isneeded. Medium-pressure steam stream (42) is sent to the steamdistribution system (11). As the gas (74) is cooled down to 200° F. (93°C.), it begins to approach the water dew-point and the condensing wateris recovered in a first knock-out drum (not shown). Subsequent coolingof gas (74) takes place against the air cooler (not shown), which usesambient air as a cooling medium, and finally the trim cooler (notshown), to achieve a final temperature of 120° F. (49° C.) using coolingwater. Ambient conditions at the location of the low-temperature gascooling system (450) will dictate the amount of air cooling and trimcooling that can be achieved. The stream leaving the trim cooler is sentto the second knock-out drum (not shown) to separate the remaining waterfrom the gas (74). The combined condensate from the knock-out drums (notshown) is sent to the ammonia recovery system (600). Dry raw gas stream(75) exits the low-temperature gas cooling system (450).

Ammonia Recovery System (600)

The low-temperature operation of the hydromethanation reactor (200)under highly reducing conditions relative to other gasificationtechnologies allows all the nitrogen released as ammonia duringdevolatilization to remain in molecular form without converting to othernitrogen oxides or decomposing to gaseous nitrogen. Ammonia can berecovered according to methods known to those skilled in the art. Aparticular embodiment of the ammonia recovery process is described next.

Referring to FIG. 4, after dry raw gas stream (75) exits low-temperaturegas cooling system (450) it is treated in an ammonia recovery system(600) to form ammonia-free dry raw gas stream (76). Ammonia recoverysystem (600) will typically comprise a series of sour water strippersand a Claus unit. Ammonia is recovered from dry raw gas stream (75) byfirst washing dry raw gas stream (75) with chilled water at 50° F. (10°C.) to remove a majority of the ammonia. The resulting ammonia scrubberbottoms liquid is combined with the condensate from the knock-out drumsand fed to a series of sour water strippers (not shown) that separatethe ammonia from liquid-phase as a primary product stream and an off-gascontaining trace amounts of ammonia, hydrogen cyanide, hydrogen sulfideand carbonyl sulfide. The off-gas stream is sent to the Claus unit (notshown) for further treatment. The ammonia recovery system (600) alsoproduces a condensate (not shown) which is recycled to hot gas scrubber(420).

The clean water leaving the sour-water strippers is devoid of dissolvedgases. A portion of this water is utilized as a liquid feed for the hotgas scrubber (420). The balance of the water is sent to the catalystrecovery system (300) as a solvent for the char washing step (notshown).

Ammonia recovery is greater than 95% of the ammonia contained in themethane-rich raw gas stream. Ammonia is typically recovered as anaqueous solution (81) of concentration 20-30 wt %. Any recovered ammoniacan be used as such or, for example, can be converted with otherby-products from the process. For example, it may be reacted withsulfuric acid to generate ammonium sulfate as a product.

Optional Booster Compressor Unit (452)

After exiting ammonia recovery system (600), ammonia-free dry raw gasstream (76) may optionally be compressed prior to treatment in acid gasremoval unit (800) to generate a compressed ammonia-free dry raw gasstream (77), as disclosed in previously incorporated US2012/0305848A1. Abooster compressor unit (452) compresses the stream to a pressure whichis higher than ammonia-free dry raw gas stream (76), and much higherthan the operating pressure of hydromethanation reactor (200).

The purpose of the booster compressor unit (452) is to take advantage ofthe higher efficiency of removal of CO₂ at higher pressures usingphysical absorbents such as refrigerated methanol. At the same time, thehydromethanation reactor (200) may be operated at a lower pressure thatis optimal for catalytic gasification and does not require the entirefront-end of the process to be designed for a higher operating pressure.Although the volume of ammonia-free dry raw gas stream (76) is higherthan that of the final methane-enriched product stream (82) from themethanation system (950), the increased power consumption of boostercompressor (452) can be partially offset by the elimination of a finalproduct compressor (952). For example, the hydromethanation reactor(200) can be operated at 500 psig (3549 kPa) and booster compressor(452) can raise the pressure of compressed ammonia-free dry raw gasstream (77) to 1050-1100 psig (7341-7686 kPa).

Booster compressor unit (452) can be a single or series of gascompressors depending on the required extent of compression, as will beunderstood by a person of ordinary skill in the art. Suitable types ofcompressors are also generally well known to those of ordinary skill inthe art, for example, compressors known to be suitable for use withsyngas streams (carbon monoxide plus hydrogen) would also be suitablefor use in connection with the present process.

As indicated above, compressed ammonia-free dry raw gas stream (77) isat a pressure higher than ammonia-free dry raw gas stream (76). In oneembodiment, the pressure of compressed ammonia-free gas raw gas stream(77) is about 20% higher or greater, or about 35% higher or greater, orabout 50% higher or greater, to about 100% higher or less, than thepressure of ammonia-free dry raw gas stream (76).

In another embodiment, the pressure of compressed ammonia-free gas rawgas stream (77) is about 720 psig (5066 kPa) or greater, or about 750psig (5272 kPa) or greater, and about 1000 psig (6996 kPa) or less, orabout 900 psig (6307 kPa) or less, or about 850 psig (5962 kPa) or less.

In another embodiment, the pressure of ammonia-free dry raw gas stream(76) is about 600 psig (4238 kPa) or less, or about 550 psig (3893 kPa)or less, or about 500 psig (3549 kPa) or less, and about 400 psig (2860kPa) or greater, or about 450 psig (3204 kPa) or greater.

Acid Gas Removal System (800)

The ammonia-free dry raw gas stream (76) leaving the ammonia recoverysystem (600) (or the compressed ammonia-free gas raw gas stream (77)leaving the optional booster compressor unit (452), if present) issubsequently fed to an acid gas removal (AGR) unit (800) to remove asubstantial portion of CO₂ and a substantial portion of the H₂S andgenerate a sweetened gas stream (80).

Acid gas removal processes typically involve contacting a gas streamwith a solvent that selectively absorbs the acid gases. Several acid gasremoval processes are commercially available and applicable fortreatment of streams (76) or (77). The main criteria for selection ofthe AGR are minimization of methane losses such that the sweetened gasstream (80) comprises at least a substantial portion (and substantiallyall) of the methane from the stream fed into acid gas removal unit(800). Typically, such losses should be about 2 mol % or less, or about1.5 mol % or less, or about 1 mol % of less, respectively, of themethane feed to the AGR.

A solvent that meets the above criteria is refrigerated methanol. Acommercially available process employing methanol as solvent is known bythe trade-name Rectisol® and is offered by Linde AG and LurgiOel-Gas-Chemie GmbH. Another commercial process that may be consideredis Selexol® (UOP LLC, Des Plaines, Ill. USA), which uses a proprietarysolvent (dimethyl ether of polyethylene glycol). Similarly, a chemicalsolvent comprised of methyldiethanolamine (MDEA) with other additivessuch as piperazine may also be used. MDEA is available from processlicensors such as BASF and Dow.

One method for removing acid gases is described in previouslyincorporated US2009/0220406A1.

At least a substantial portion (e.g., substantially all) of the CO₂and/or H₂S (and other remaining trace contaminants) should be removedvia the acid gas removal processes. “Substantial” removal in the contextof acid gas removal means removal of a high enough percentage of thecomponent such that a desired end product can be generated. The actualamounts of removal may thus vary from component to component. For“pipeline-quality natural gas”, only trace amounts (at most) of H₂S canbe present, although higher (but still small) amounts of CO₂ may betolerable.

The resulting sweetened gas stream (80) will generally comprise CH₄, H₂and CO (for the downstream methanation), and typically traces of H₂O.

Any recovered H₂S (78) from the acid gas removal (and other processessuch as sour water stripping) can be converted to elemental sulfur byany method known to those skilled in the art, including the Clausprocess. Sulfur can be recovered as a molten liquid.

Any recovered CO₂ (79) from the acid gas removal can be compressed fortransport in CO₂ pipelines, industrial use, and/or sequestration forstorage or other processes such as enhanced oil recovery and can also beused for other process operations (such as in certain aspects catalystrecovery and feed preparation).

The resulting sweetened gas stream (80) can be further processed asdescribed below to produce pipeline quality SNG.

Methanation System (950)

The sweetened gas stream (80) is fed to a methanation system (950) togenerate additional methane from the carbon monoxide and hydrogen thatmay be present in those streams, resulting in a methane-enriched productstream (82).

The methanation system (950) will typically comprise one or moremethanation reactors or catalytic methanators (not shown), e.g., asingle-stage methanation reactor, a series of single-stage methanationreactors or a multistage reactor. Methanation reactors include, withoutlimitation, fixed bed, moving bed or fluidized bed reactors. See, forinstance, U.S. Pat. Nos. 3,958,957, 4,252,771, 3,996,014 and 4,235,044.Methanation reactors and catalysts are generally commercially available.The catalyst used in the methanation, and methanation conditions, aregenerally known to those of ordinary skill in the relevant art, and willdepend, for example, on the temperature, pressure, flow rate andcomposition of the incoming gas stream.

The methanation system (950) will also typically comprise a heatrecovery system with heat exchanger units. As the methanation reactionis highly exothermic, the heat recovery system generates high-pressure,superheated or saturated steam in the heat exchanger units to remove aportion of the heat energy. The recovered heat energy is utilized togenerate a high-pressure steam stream (43) from boiler feed water (39d).

A final product compressor (952) raises the pressure of themethane-enriched product stream (82) from the methanation system (950)to the required pipeline pressure to form a pipeline quality SNG (83).Alternatively, the methane-enriched product stream (82) can be furtherprocessed, when necessary, to separate and recover CH₄ by any suitablegas separation method known to those skilled in the art including, butnot limited to, cryogenic distillation and the use of molecular sievesor gas separation (e.g., ceramic) membranes. Additional gas purificationmethods include, for example, the generation of methane hydrate asdisclosed in previously incorporated US2009/0260287A1, US2009/0259080A1and US2009/0246120A1.

Pipeline Quality Natural Gas

The processes and systems described are capable of generating“pipeline-quality natural gas” (or “pipeline-quality substitute naturalgas”) from the hydromethanation of non-gaseous carbonaceous materials. A“pipeline-quality natural gas” typically refers to a methane-containingstream that meets the conditions in Table 1:

TABLE 1 Parameter Value Units Source Higher Heating Value >35.367 MJ/Nm³1 (HHV) >950 BTU/scf 1 Wobbe Index (WI) >46.1-56.5 MJ/Nm³ 2 RelativeDensity   0.55-0.70 — 2 Sulfur Content <0.0023 kg/100 Nm³ 1 <0.00014lb/100 scf 1 CO₂ Content <2 vol. % 1 Moisture Content <62.5 kg/1000 1Nm³ <3.9 lb/1000 scf 1 Temperature <49 (120) ° C. (° F.) 1Toxic/Corrosive None — 1 contaminants

The specification for the product quality of SNG in terms of gascomposition is listed in Table 2:

TABLE 2 Component Value Units CH₄ 94-98 mol % H₂ <2 mol % CO <100 ppmvCO₂ <97 ppmv

In the process described, the pipeline quality SNG (83) satisfies thespecification of Tables 1 and 2.

Steam Generation and Distribution System

The hydromethanation process requires steam at several differentpressures. First, steam is needed as a reactant in the hydromethanationreactor (200). Steam is fed to the hydromethanation reactor (200) at apressure that is higher than the reactor pressure by at least 50 psig(446 kPa). Although the reactor can be operated with saturated steam, anenergy penalty in terms of increased oxygen use, decreased methaneproduction and increased carbon dioxide production must be incurred. Asa result, superheated steam at 510° C. (950° F.) at the requiredpressure is preferred to maximize the overall process thermalefficiency. Second, steam is required as a utility to perform variousheating duties such as evaporation/crystallization of catalyst solution,reboiler for the AGR and ammonia recovery system, etc.

Referring to FIG. 4, the steam distribution system (11) receives thesteam generated by various sources and distributes them to consumerswithin the process. The main process heat exchanger or boiler (400)following the hydromethanation reactor (200) and the heat exchangerunits within the methanation system (950) produce high pressure steam ofthe required quality for the hydromethanation reactor (200). Asdiscussed previously, the temperature of the steam is normally maximizedto improve efficiency. The hydromethanation reactor (200) steamrequirements, supplied by steam stream (12), are met by distributinghigh-pressure steam streams (40 and 43), via steam distribution system(11). The high-pressure, saturated or superheated steam in excess of thehydromethanation reactor (200) requirements is let down in pressure to alevel of 50 psig (446 kPa). The saturated steam from the intermediatepressure boiler (410) at 150 psig (1136 kPa) is also let down inpressure to a level of 50 psig (446 kPa). The low-temperature gascooling system (450) also produces medium-pressure (50 psig) steam (42)by recovery of lower grade heat. All sources of 50 psig steam serve as aheat-transfer media for various consumers within the process. Excess 50psig (446 kPa) steam is let down to 30 psig (304 kPa) and combines withsources of 30 psig (304 kPa) steam within the low temperature gascooling to be distributed to various consumers within the process. Thevarious steam sources produce sufficient steam at the required levels tomeet the requirements of various consumers. As a result, the overallprocess is steam balanced. Any high-pressure steam in excess of processrequirements may be converted to power as described below. The processhas a steam demand and a power demand that are met by internal energyintegration such that the process requires no net import of steam orpower.

Power Generation System

As shown in FIG. 2, a gas turbine combined cycle system, comprising anair compressor (991), a high-pressure gas turbine combustion chamber(992), a gas expander (993) and a heat recovery steam generator (975),is employed to generate electric power from the low or medium BTUoxidation reactor product gas (88) exiting the oxidation reactor (380)operated at a high pressure. The gas turbine combined cycle systemgenerates power by burning gas (88) with atmospheric air (90) atelevated pressure (typically about 10 bar) in the combustion chamber(992) and expanding the combustion gases through the gas expander (993).As the gas expands and cools, it drives the gas expander which in turnwill power the air compressor (991) and electric generator (99 a). Theamount of gross power generated depends on the pressure, temperature,volume and molecular weight of the combustion gases and the efficiencyof the gas expander. The net power generated is calculated bysubtracting the power needed to compress the air and the efficiencylosses of the electric generator from the gross power generated by thegas expander.

If the oxidation reactor product gas (88) is pressurized but has nochemical heating value since it is comprised of CO₂, steam and nitrogen,some work and electric power can still be recovered by expanding the hotoxidation reactor product gas (88) through the gas expander (993)coupled to an electric generator (99 a) as shown in FIG. 3.

The expander exhaust gas (97) contains sufficient thermal energy toproduce steam and subsequently more power. The expander exhaust gas (97)typically is above 454° C. (850° F.) and may be used to generatesuperheated steam in the heat recovery steam generator (975). The heatrecovery steam generator (975) is an insulated low-pressure vessel withan arrangement of tubes and coils that are designed to efficiently coolthe exhaust gas and convert boiler feed water into steam. Sometimes theheat recovery steam generator (975) is equipped with burners to providesupplemental heat in order to increase the temperature and/or pressureof the steam.

The steam produced in the heat recovery steam generator (975) isdirected to a steam-turbine generator which comprises a steam turbine(980) that drives an electric generator (99 b). The steam turbine mayinclude any number of stages whereby some steam may be extracted atdifferent pressures and temperatures, if desired. In addition, theextracted steam can be re-heated and injected into a subsequent stage toimprove the net power generated by the steam turbine.

The electrical generators can be of any kind, AC or DC, that are widelyused today. All electrical generators fundamentally consist of twoparts: (1) electro-magnets or permanent magnets that produce a magneticflux and (2) a laminated steel core that carrying conductors. Therelative motion between the conductors and magnetic field produces avoltage in accordance to Faraday's law.

Water Treatment and Recovery

Residual contaminants in waste water resulting from any one or more ofthe trace contaminant removal, sour shift, ammonia removal, acid gasremoval and/or catalyst recovery processes can be removed in a wastewater treatment unit (not shown) to allow recycling of the recoveredwater within the plant and/or disposal of the water from the plantprocess according to any methods known to those skilled in the art.Depending on the feedstock and reaction conditions, such residualcontaminants can comprise, for example, aromatics, CO, CO₂, H₂S, COS,HCN, NH₃, and Hg. For example, H₂S and HCN can be removed byacidification of the waste water to a pH of about 3, treating the acidicwaste water with an inert gas in a stripping column, and increasing thepH to about 10 and treating the waste water a second time with an inertgas to remove ammonia (see U.S. Pat. No. 5,236,557). H₂S can be removedby treating the waste water with an oxidant in the presence of residualcoke particles to convert the H₂S to insoluble sulfates which may beremoved by flotation or filtration (see U.S. Pat. No. 4,478,725).Aromatics can be removed by contacting the waste water with acarbonaceous char optionally containing mono- and divalent basicinorganic compounds (e.g., the solid char product or the depleted charafter catalyst recovery, supra) and adjusting the pH (see U.S. Pat. No.4,113,615). Trace amounts of aromatics (C₆H₆, C₇H₈, C₁₀H₈) can also beremoved by extraction with an organic solvent followed by treatment ofthe waste water in a stripping column (see U.S. Pat. Nos. 3,972,693,4,025,423 and 4,162,902).

Multi-Train Processes

Each process may be performed in one or more processing units. Forexample, one or more hydromethanation reactors may be supplied with thecarbonaceous feedstock from one or more catalyst loading and/orfeedstock preparation unit operations. Similarly, the methane-enrichedraw product streams generated by one or more hydromethanation reactorsmay be processed or purified separately or via their combination atvarious downstream points depending on the particular systemconfiguration, as discussed, for example, in previously incorporatedUS2009/0324458A1, US2009/0324459A1, US2009/0324460A1, US2009/0324461A1and US2009/0324462A1.

In certain embodiments, the processes utilize two or morehydromethanation reactors (e.g., 2-4 hydromethanation reactors). In suchembodiments, the processes may contain divergent processing units (i.e.,less than the total number of hydromethanation reactors) prior to thehydromethanation reactors for ultimately providing the catalyzedcarbonaceous feedstock to the plurality of hydromethanation reactors,and/or convergent processing units (i.e., less than the total number ofhydromethanation reactors) following the hydromethanation reactors forprocessing the plurality of methane-enriched raw product streamsgenerated by the plurality of hydromethanation reactors.

When the systems contain convergent processing units, each of theconvergent processing units can be selected to have a capacity to acceptgreater than a 1/n portion of the total feed stream to the convergentprocessing units, where n is the number of convergent processing units.Similarly, when the systems contain divergent processing units, each ofthe divergent processing units can be selected to have a capacity toaccept greater than a 1/m portion of the total feed stream supplying theconvergent processing units, where m is the number of divergentprocessing units.

While a number of example embodiments have been provided, the variousaspects and embodiments disclosed herein are for purposes ofillustration and are not intended to be limiting. Other embodiments canbe used, and other changes can be made, without departing from thespirit and scope of the subject matter presented herein. It will bereadily understood that the aspects of the disclosure, as generallydescribed herein, and illustrated in the figures, can be arranged,substituted, combined, separated, and designed in a wide variety ofdifferent configurations, all of which are explicitly contemplatedherein.

We claim:
 1. A process for generating, from a non-gaseous carbonaceousmaterial and a hydromethanation catalyst, (1) a fines-cleanedmethane-enriched raw product gas stream, (2) an oxidation reactorproduct gas and (3) an oxidation reactor ash product stream, the processcomprising the steps of: a) preparing a carbonaceous feedstock from thenon-gaseous carbonaceous material; b) introducing the carbonaceousfeedstock, the hydromethanation catalyst, steam and oxygen into ahydromethanation reactor, the hydromethanation reactor comprising afluidized bed, a disengagement zone above the fluidized bed, and a gasmixing zone below the fluidized bed; c) reacting the carbonaceousfeedstock in the hydromethanation reactor in the presence of carbonmonoxide, hydrogen, steam and hydromethanation catalyst, and at anoperating temperature from about 400° F. (about 205° C.) up to about1500° F. (about 816° C.), and an operating pressure of at least about250 psig (about 1825 kPa), to produce a methane-enriched raw productgas, heat energy and a by-product char; d) withdrawing a stream ofmethane-enriched raw product gas from the hydromethanation reactor asthe methane-enriched raw product gas stream, wherein themethane-enriched raw product gas stream comprises methane, carbonmonoxide, hydrogen, carbon dioxide, hydrogen sulfide, steam, andentrained solids; e) removing a substantial portion of the entrainedsolids from the methane-enriched raw product gas stream to generate asolids-depleted, methane-enriched raw product gas stream and a recoveredprimary solids stream; f) removing a substantial portion of any finesfrom the solids-depleted, methane-enriched raw product gas stream togenerate the fines-cleaned methane-enriched raw product gas stream and arecovered secondary fines stream; g) withdrawing a stream of by-productchar from the hydromethanation reactor as the hydromethanation reactorchar product stream, wherein the hydromethanation reactor char productstream comprises a carbon content and entrained hydromethanationcatalyst; h) extracting a portion of the entrained catalyst from thehydromethanation reactor char product stream by the steps of: (i)feeding all or a portion of the hydromethanation reactor char productstream to a catalyst recovery unit; (ii) withdrawing a stream ofcatalyst-depleted char from the catalyst recovery unit as the washedchar product stream, the washed char product stream comprising a carboncontent; and (iii) withdrawing a stream of liberated hydromethanationcatalyst from the catalyst recovery unit as a recovered hydromethanationcatalyst stream; i) feeding into an oxidation reactor (1) all or aportion of the washed char product stream, and (2) an oxygen-containinggas stream; j) reacting at least a portion of the carbon content of thewashed char product stream with the oxygen-containing gas stream in theoxidation reactor to produce (1) an oxidation reactor ash productstream, and (2) an oxidation reactor product gas comprising steam andcarbon dioxide; and k) withdrawing the oxidation reactor product gasfrom the oxidation reactor.
 2. The process of claim 1, furthercomprising the steps of: (A) operating the oxidation reactor at apressure greater than 791 kPa (100 psig); (B) (i) when the oxidationreactor product gas further comprises carbon monoxide and hydrogen,feeding all or a portion of the oxidation reactor product gas into a gasturbine combined cycle system to generate power and steam; or (ii) whencarbon monoxide and hydrogen are not present in the oxidation reactorproduct gas, feeding all or a portion of the oxidation reactor productgas into a gas expander and a heat recovery steam generator to generatepower and steam; and (C) cooling or quenching the oxidation reactor ashproduct stream.
 3. The process of claim 1, further comprising the stepsof: (A) operating the oxidation reactor at a pressure less than 400 kPa(44 psig); (B) (i) when the oxidation reactor product gas furthercomprises carbon monoxide and hydrogen, feeding all or a portion of theoxidation reactor product gas to a heat recovery steam generatorequipped with an auxiliary burner to generate steam; or (ii) when carbonmonoxide and hydrogen are not present in the oxidation reactor productgas, feeding all or a portion of the oxidation reactor product gas to aheat recovery steam generator to generate steam; and (C) cooling orquenching the oxidation reactor ash product stream.
 4. The process ofclaim 2, further comprising the steps of (D) converting all or a portionof the steam generated in step (B) to power in a steam-turbinegenerator, and (E) exporting any unused steam.
 5. The process of claim3, further comprising the steps of (D) converting all or a portion ofthe steam generated in step (B) to power in a steam-turbine generator,and (E) exporting any unused steam.
 6. The process of claim 1, whereinthe operating pressure in the hydromethanation reactor ranges from atleast about 250 psig (about 1825 kPa) up to about 1200 psig (about 8375kPa).
 7. The process of claim 1, wherein the oxidation reactor is afluidized-bed oxidation reactor.
 8. The process of claim 1, wherein thehydromethanation catalyst comprises an alkali metal.
 9. The process ofclaim 8, wherein the alkali metal is potassium.
 10. The process of claim1, further comprising the steps of: (f1) feeding the recovered primarysolids stream into the hydromethanation reactor; and (f2) feeding all ora portion of the recovered secondary fines stream into the catalystrecovery unit.
 11. The process of claim 1, wherein the hydromethanationcatalyst comprises at least a portion of the recovered hydromethanationcatalyst stream.
 12. The process of claim 1, further comprising thesteps of: l) introducing the fines-cleaned methane-enriched raw productgas stream into a heat exchanger unit to remove heat energy and generatea cooled methane-enriched raw product stream; m) steam shifting aportion of the carbon monoxide in the cooled methane-enriched rawproduct stream in a shift reactor system to generate a hydrogen-enrichedraw product gas stream with a molar ratio of hydrogen to carbon monoxideof close to 3; n) dehydrating the hydrogen-enriched raw product gasstream in a low-temperature gas cooling system, to generate a dry rawgas stream, the dry raw gas stream comprising carbon dioxide, hydrogensulfide, hydrogen, carbon monoxide and methane; and o) removing asubstantial portion of the carbon dioxide and a substantial portion ofthe hydrogen sulfide from the dry raw gas stream in an acid gas removalunit to produce a sweetened gas stream comprising a substantial portionof the hydrogen, carbon monoxide and methane from the dry raw gasstream.
 13. The process of claim 12, further comprising the steps of: p)reacting the carbon monoxide and hydrogen in the sweetened gas stream ina methanation system in the presence of a methanation catalyst toproduce heat energy and a pipeline quality substitute natural gasstream; q) recovering the heat energy from the catalytic methanationsystem; and r) utilizing at least a portion of the recovered heat energyto generate and superheat a steam stream.
 14. The process of claim 13,wherein the process has a steam demand and a power demand that are metby internal energy integration such that the process requires no netimport of steam or power.